Optimized liquid-phase oxidation

ABSTRACT

Disclosed is an optimized process and apparatus for more efficiently and economically carrying out the liquid-phase oxidation of an oxidizable compound. Such liquid-phase oxidation is carried out in a bubble column reactor that provides for a highly efficient reaction at relatively low temperatures. When the oxidized compound is para-xylene and the product from the oxidation reaction is crude terephthalic acid (CTA), such CTA product can be purified and separated by more economical techniques than could be employed if the CTA were formed by a conventional high-temperature oxidation process.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims priority to U.S. Provisional Application SerialNo. 60/606,618, filed Sep. 2, 2004 and 60/631,326, filed Nov. 29, 2004,the disclosures of which are incorporated herein by reference in theirentirety.

FIELD OF THE INVENTION

This invention relates generally to a process for the liquid-phase,catalytic oxidation of an aromatic compound. One aspect of the inventionconcerns the partial oxidation of a dialkyl aromatic compound (e.g.,para-xylene) to produce a crude aromatic dicarboxylic acid (e.g., crudeterephthalic acid), which can thereafter be subjected to purificationand separation. Another aspect of the invention concerns an improvedbubble column reactor that provides for a more effective and economicalliquid-phase oxidation process.

BACKGROUND OF THE INVENTION

Liquid-phase oxidation reactions are employed in a variety of existingcommercial processes. For example, liquid-phase oxidation is currentlyused for the oxidation of aldehydes to acids (e.g., propionaldehyde topropionic acid), the oxidation of cyclohexane to adipic acid, and theoxidation of alkyl aromatics to alcohols, acids, or diacids. Aparticularly significant commercial oxidation process in the lattercategory (oxidation of alkyl aromatics) is the liquid-phase catalyticpartial oxidation of para-xylene to terephthalic acid. Terephthalic acidis an important compound with a variety of applications. The primary useof terephthalic acid is as a feedstock in the production of polyethyleneterephthalate (PET). PET is a well-known plastic used in greatquantities around the world to make products such as bottles, fibers,and packaging.

In a typical liquid-phase oxidation process, including partial oxidationof para-xylene to terephthalic acid, a liquid-phase feed stream and agas-phase oxidant stream are introduced into a reactor and form amulti-phase reaction medium in the reactor. The liquid-phase feed streamintroduced into the reactor contains at least one oxidizable organiccompound (e.g., para-xylene), while the gas-phase oxidant streamcontains molecular oxygen. At least a portion of the molecular oxygenintroduced into the reactor as a gas dissolves into the liquid phase ofthe reaction medium to provide oxygen availability for the liquid-phasereaction. If the liquid phase of the multi-phase reaction mediumcontains an insufficient concentration of molecular oxygen (i.e., ifcertain portions of the reaction medium are “oxygen-starved”),undesirable side-reactions can generate impurities and/or the intendedreactions can be retarded in rate. If the liquid phase of the reactionmedium contains too little of the oxidizable compound, the rate ofreaction may be undesirably slow. Further, if the liquid phase of thereaction medium contains an excess concentration of the oxidizablecompound, additional undesirable side-reactions can generate impurities.

Conventional liquid-phase oxidation reactors are equipped with agitationmeans for mixing the multi-phase reaction medium contained therein.Agitation of the reaction medium is supplied in an effort to promotedissolution of molecular oxygen into the liquid phase of the reactionmedium, maintain relatively uniform concentrations of dissolved oxygenin the liquid phase of the reaction medium, and maintain relativelyuniform concentrations of the oxidizable organic compound in the liquidphase of the reaction medium.

Agitation of the reaction medium undergoing liquid-phase oxidation isfrequently provided by mechanical agitation means in vessels such as,for example, continuous stirred tank reactors (CSTRs). Although CSTRscan provide thorough mixing of the reaction medium, CSTRs have a numberof drawbacks. For example, CSTRs have a relatively high capital cost dueto their requirement for expensive motors, fluid-sealed bearings anddrive shafts, and/or complex stirring mechanisms. Further, the rotatingand/or oscillating mechanical components of conventional CSTRs requireregular maintenance. The labor and shutdown time associated with suchmaintenance adds to the operating cost of CSTRs. However, even withregular maintenance, the mechanical agitation systems employed in CSTRsare prone to mechanical failure and may require replacement overrelatively short periods of time.

Bubble column reactors provide an attractive alternative to CSTRs andother mechanically agitated oxidation reactors. Bubble column reactorsprovide agitation of the reaction medium without requiring expensive andunreliable mechanical equipment. Bubble column reactors typicallyinclude an elongated upright reaction zone within which the reactionmedium is contained. Agitation of the reaction medium in the reactionzone is provided primarily by the natural buoyancy of gas bubbles risingthrough the liquid phase of the reaction medium. This natural-buoyancyagitation provided in bubble column reactors reduces capital andmaintenance costs relative to mechanically agitated reactors. Further,the substantial absence of moving mechanical parts associated withbubble column reactors provides an oxidation system that is less proneto mechanical failure than mechanically agitated reactors.

When liquid-phase partial oxidation of para-xylene is carried out in aconventional oxidation reactor (CSTR or bubble column), the productwithdrawn from the reactor is typically a slurry comprising crudeterephthalic acid (CTA) and a mother liquor. CTA contains relativelyhigh levels of impurities (e.g., 4-carboxybenzaldehyde, para-toluicacid, fluorenones, and other color bodies) that render it unsuitable asa feedstock for the production of PET. Thus, the CTA produced inconventional oxidation reactors is typically subjected to a purificationprocess that converts the CTA into purified terephthalic acid (PTA)suitable for making PET.

One typical purification process for converting CTA to PTA includes thefollowing steps: (1) replacing the mother liquor of the CTA-containingslurry with water, (2) heating the CTA/water slurry to dissolve the CTAin water, (3) catalytically hydrogenating the CTA/water solution toconvert impurities to more desirable and/or easily-separable compounds,(4) precipitating the resulting PTA from the hydrogenated solution viamultiple crystallization steps, and (5) separating the crystallized PTAfrom the remaining liquids. Although effective, this type ofconventional purification process can be very expensive. Individualfactors contributing to the high cost of conventional CTA purificationmethods include, for example, the heat energy required to promotedissolution of the CTA in water, the catalyst required forhydrogenation, the hydrogen stream required for hydrogenation, the yieldloss caused by hydrogenation of some terephthalic acid, and the multiplevessels required for multi-step crystallization. Thus, it would bedesirable to provide a CTA product that could be purified withoutrequiring heat-promoted dissolution in water, hydrogenation, and/ormulti-step crystallization.

OBJECTS OF THE INVENTION

It is, therefore, an object of the present invention to provide a moreeffective and economical liquid-phase oxidation reactor and process.

Another object of the invention is to provide a more effective andeconomical reactor and process for the liquid-phase catalytic partialoxidation of para-xylene to terephthalic acid.

Still another object of the invention is to provide a bubble columnreactor that facilitates improved liquid-phase oxidation reactions withreduced formation of impurities.

Yet another object of the invention is to provide a more effective andeconomical system for producing pure terephthalic acid (PTA) vialiquid-phase oxidation of para-xylene to produce crude terephthalic acid(CTA) and subsequently, purifying the CTA to PTA.

A further object of the invention is to provide a bubble column reactorfor oxidizing para-xylene and producing a CTA product capable of beingpurified without requiring heat-promoted dissolution of the CTA inwater, hydrogenation of the dissolved CTA, and/or multi-stepcrystallization of the hydrogenated PTA.

It should be noted that the scope of the present invention, as definedin the appended claims, is not limited to processes or apparatusescapable of realizing all of the objects listed above. Rather, the scopeof the claimed invention may encompass a variety of systems that do notaccomplish all or any of the above-listed objects. Additional objectsand advantages of the present invention will be readily apparent to oneskilled in the art upon reviewing the following detailed description andassociated drawings.

SUMMARY OF THE INVENTION

One embodiment of the present invention concerns a process comprisingthe following steps: (a) forming a three-phase medium in a reaction zoneof a bubble column reactor; (b) introducing at least a portion of thethree-phase medium into a deaeration vessel having one or more generallyupright sidewalls defining a deaeration zone, wherein the maximumhorizontal cross-sectional area of the deaeration zone is less thanabout 25 percent of the maximum horizontal cross-sectional area of thereaction zone; and (c) separating at least a portion of the gas phase ofthe three-phase medium from the solid and liquid phases of the medium inthe deaeration zone to thereby form a substantially deaerated slurry,wherein the separating is caused primarily by the natural buoyancy ofthe gas phase in the solid and liquid phases.

Another embodiment of the present invention concerns a processcomprising the following steps: (a) oxidizing an oxidizable compound ina liquid phase of a three-phase reaction medium contained in a bubblecolumn reactor, wherein the bubble column reactor includes a vesselshell defining an outlet; (b) introducing at least a portion of thethree-phase reaction medium into a deaeration vessel comprising one ormore generally upright sidewalls at least partly disposed within thevessel shell, wherein the upright sidewalls define an internaldeaeration zone that is in fluid flow communication with the outlet; and(c) separating at least a substantial portion of the gas phase of thethree-phase reaction medium from the solid and liquid phases of thereaction medium in the deaeration zone to thereby form a substantiallydeaerated slurry, wherein the separating is caused primarily by thenatural buoyancy of the gas phase in the solid and liquid phases.

Still another embodiment of the present invention concerns a processcomprising the following steps: (a) oxidizing an oxidizable compound ina liquid phase of a three-phase reaction medium contained in a reactionzone of a bubble column reactor, wherein the bubble column reactorincludes a vessel sidewall defining a side outlet that is elevated fromthe lower end of the reaction zone; (b) withdrawing at least a portionof the three-phase reaction medium from the reaction zone via the sideoutlet; (c) introducing at least a portion of the withdrawn three-phasereaction medium into a deaeration vessel comprising one or moregenerally upright sidewalls defining a deaeration zone; and (d)separating at least a substantial portion of the gas phase of thethree-phase reaction medium from the solid and liquid phases of thereaction medium in the deaeration zone to thereby form a substantiallydeaerated slurry, wherein the separating is caused primarily by thenatural buoyancy of the gas phase in the solid and liquid phases.

Yet another embodiment of the present invention concerns a process formaking purified terephthalic acid comprising the following steps: (a)oxidizing para-xylene in a liquid phase of a three-phase reaction mediumcontained in a reaction zone of a bubble column reactor to thereby formcrude terephthalic acid; (b) introducing at least a portion of thethree-phase reaction medium into a deaeration vessel having one or moregenerally upright sidewalls defining a deaeration zone, wherein themaximum horizontal cross-sectional area of the deaeration zone is lessthan about 25 percent of the maximum horizontal cross-sectional area ofthe reaction zone; (c) separating at least a substantial portion of thegas phase of the three-phase medium from the solid and liquid phases ofthe medium in the deaeration zone to thereby form a substantiallydeaerated slurry comprising the crude terephthalic acid, wherein theseparating is caused primarily by the natural buoyancy of the gas phasein the solid and liquid phases; and (d) oxidizing at least a portion ofthe crude terephthalic acid from the deaerated slurry to thereby formpurer terephthalic acid.

A further embodiment of the present invention concerns a bubble columnreactor for contacting a predominately liquid-phase stream and apredominately gas-phase stream. The bubble column reactor includes avessel shell, a deaeration vessel, and a purification system. The vesselshell defines a reaction zone and a vessel outlet. The deaeration vesselis coupled in fluid flow communication with the vessel outlet. Thedeaeration vessel presents an upper inlet and a lower outlet. Thedeaeration vessel comprises one or more upright sidewalls defining aninternal deaeration zone in fluid flow communication with the upperinlet and the lower outlet. The maximum horizontal cross-sectional areaof the deaeration zone is less than about 25 percent of the maximumhorizontal cross-sectional area of the reaction zone.

A still further embodiment of the present invention concerns a systemfor producing and purifying terephthalic acid. The system includes areactor and a deaeration vessel. The reactor facilitates the productionof crude terephthalic acid in a three-phase reaction medium. The reactordefines a reactor outlet for discharging at least the solid and liquidphases of the reaction medium from the reactor. The deaeration vesselfacilitates the disengagement of the gas phase of the three-phasereaction medium from the solid and liquid phases of the three-phasereaction medium to thereby provide a substantially deaerated slurry. Thedeaeration vessel is coupled in fluid flow communication with thereactor outlet. The deaeration vessel defines a reaction medium inletfor receiving the three-phase reaction medium and a slurry outlet fordischarging the deaerated slurry. The slurry outlet is located lowerthan the reaction medium inlet. The deaeration vessel comprises one ormore upright sidewalls defining an internal deaeration zone verticallydisposed between the reaction medium inlet and the slurry outlet. Thedeaeration zone has a maximum height in the range of from about 0.3 toabout 5 meters and a maximum horizontal cross-sectional area in therange of from about 0.02 to about 3 meters squared. The purificationsystem receives the deaerated slurry from the slurry outlet and purifiesthe deaerated slurry.

BRIEF DESCRIPTION OF THE DRAWINGS

Preferred embodiments of the invention are described in detail belowwith reference to the attached drawing figures, wherein;

FIG. 1 is a side view of an oxidation reactor constructed in accordancewith one embodiment of the present invention, particularly illustratingthe introduction of feed, oxidant, and reflux streams into the reactor,the presence of a multi-phase reaction medium in the reactor, and thewithdrawal of a gas and a slurry from the top and bottom of the reactor,respectively;

FIG. 2 is an enlarged sectional side view of the bottom of the bubblecolumn reactor taken along line 2-2 in FIG. 3, particularly illustratingthe location and configuration of a oxidant sparger used to introducethe oxidant stream into the reactor;

FIG. 3 is a top view of the oxidant sparger of FIG. 2, particularlyillustrating the oxidant openings in the top of the oxidant sparger;

FIG. 4 is a bottom view of the oxidant sparger of FIG. 2, particularlyillustrating the oxidant openings in the bottom of the oxidant sparger;

FIG. 5 is a sectional side view of the oxidant sparger taken along line5-5 in FIG. 3, particularly illustrating the orientation of the oxidantopenings in the top and bottom of the oxidant sparger;

FIG. 6 is an enlarged side view of the bottom portion of the bubblecolumn reactor, particular illustrating a system for introducing thefeed stream into the reactor at multiple, vertically-space locations;

FIG. 7 is a sectional top view taken along line 7-7 in FIG. 6,particularly illustrating how the feed introduction system shown in FIG.6 distributes the feed stream into in a preferred radial feed zone (FZ)and more than one azimuthal quadrant (Q₁, Q₂, Q₃, Q₄);

FIG. 8 is a sectional top view similar to FIG. 7, but illustrating analternative means for discharging the feed stream into the reactor usingbayonet tubes each having a plurality of small feed openings;

FIG. 9 is an isometric view of an alternative system for introducing thefeed stream into the reaction zone at multiple vertically-spacelocations without requiring multiple vessel penetrations, particularlyillustrating that the feed distribution system can be at least partlysupported on the oxidant sparger;

FIG. 10 is a side view of the single-penetration feed distributionsystem and oxidant sparger illustrated in FIG. 9;

FIG. 11 is a sectional top view taken along line 11-11 in FIG. 10 andfurther illustrating the single-penetration feed distribution systemsupported on the oxidant sparger;

FIG. 12 is an isometric view of an alternative oxidant sparger havingall of the oxidant openings located in the bottom of the ring member;

FIG. 13 is a top view of the alternative oxidant sparger of FIG. 12;

FIG. 14 is a bottom view of the alternative oxidant sparger of FIG. 12,particularly illustrating the location of the bottom openings forintroducing the oxidant stream into the reaction zone;

FIG. 15 is a sectional side view of the oxidant sparger taken along line15-15 in FIG. 13, particularly illustrating the orientation of the loweroxidant openings;

FIG. 16 is a side view of a bubble column reactor equipped with aninternal deaeration vessel near the bottom outlet of the reactor;

FIG. 17 is an enlarged sectional side view of the lower portion of thebubble column reactor of FIG. 16 taken along line 17-17 in FIG. 18,particularly illustrating the configuration of the internal deaerationvessel positioned at the bottom outlet of the bubble column reactor;

FIG. 18 is a sectional top view taken along line 18-18 in FIG. 16,particularly illustrating a vortex breaker disposed in the deaerationvessel;

FIG. 19 is a side view of a bubble column reactor equipped with anexternal deaeration vessel and illustrating the manner in which aportion of the deaerated slurry exiting the bottom of the deaerationvessel can be used to flush out a de-inventorying line coupled to thebottom of the reactor;

FIG. 20 is a side view of a bubble column reactor equipped with a hybridinternal/external deaeration vessel for disengaging the gas phase of areaction medium withdrawn from an elevated side location in the reactor;

FIG. 21 is a side view of a bubble column reactor equipped with analternative hybrid deaeration vessel near the bottom of the reactor;

FIG. 22 is an enlarged sectional side view of the lower portion of thebubble column reactor of FIG. 21, particularly illustrating the use ofan alternative oxidant sparger employing inlet conduits that receive theoxidant stream through the bottom head of the reactor;

FIG. 23 is an enlarged sectional side view similar to FIG. 22,particularly illustrating an alternative means for introducing theoxidant stream into the reactor via a plurality of openings in the lowerhead of the reactor and, optionally, employing impingement plates tomore evenly distribute the oxidant stream in the reactor;

FIG. 24 is a side view of a bubble column reactor employing an internalflow conduit to help improve dispersion of an oxidizable compound byrecirculating a portion of the reaction medium from an upper portion ofthe reactor to a lower portion of the reactor;

FIG. 25 is a side view of a bubble column reactor employing an externalflow conduit to help improve dispersion of the oxidizable compound byrecirculating a portion of the reaction medium from an upper portion ofthe reactor to a lower portion of the reactor;

FIG. 26 is a sectional side view of a horizontal eductor that can beused to improve dispersion of the oxidizable compound in an oxidationreactor, particularly illustrating an eductor that uses incoming liquidfeed to draw reaction medium into the eductor and discharges the mixtureof feed and reaction medium into a reaction zone at high velocity;

FIG. 27 is a sectional side view of a vertical eductor that can be usedimprove dispersion of the oxidizable compound in an oxidation reactor,particularly illustrating an eductor that combines the liquid feed andinlet gas and uses the combined two-phase fluid to draw reaction mediuminto the eductor and discharge the mixture of liquid feed, inlet gas,and reaction medium into a reaction zone at high velocity;

FIG. 28 is a side view of a bubble column reactor containing amulti-phase reaction medium, particularly illustrating the reactionmedium being theoretically partitioned into 30 horizontal slices ofequal volume in order to quantify certain gradients in the reactionmedium;

FIG. 29 is a side view of a bubble column reactor containing amulti-phase reaction medium, particularly illustrating first and seconddiscrete 20-percent continuous volumes of the reaction medium that havesubstantially different oxygen concentrations and/or oxygen consumptionrates;

FIG. 30 is a side view of two stacked reaction vessels, with or withoutoptional mechanical agitation, containing a multi-phase reaction medium,particularly illustrating that the vessels contain discrete 20-percentcontinuous volumes of the reaction medium having substantially differentoxygen concentrations and/or oxygen consumption rates;

FIG. 31 is a side view of three side-by-side reaction vessels, with orwithout optional mechanical agitation, containing a multi-phase reactionmedium, particularly illustrating that the vessels contain discrete20-percent continuous volumes of the reaction medium havingsubstantially different oxygen concentrations and/or oxygen consumptionrates;

FIGS. 32A and 32B are magnified views of crude terephthalic acid (CTA)particles produced in accordance with one embodiment of the presentinvention, particularly illustrating that each CTA particle is a lowdensity, high surface area particle composed of a plurality ofloosely-bound CTA sub-particles;

FIG. 33A and 33B are magnified views of a conventionally-produced CTA,particularly illustrating that the conventional CTA particle has alarger particle size, lower density, and lower surface area than theinventive CTA particle of FIGS. 32A and 32B;

FIG. 34 is a simplified process flow diagram of a prior art process formaking purified terephthalic acid (PTA); and

FIG. 35 is a simplified process flow diagram of a process for making PTAin accordance with one embodiment of the present invention.

DETAILED DESCRIPTION

One embodiment of the present invention concerns the liquid-phasepartial oxidation of an oxidizable compound. Such oxidation ispreferably carried out in the liquid phase of a multi-phase reactionmedium contained in one or more agitated reactors. Suitable agitatedreactors include, for example, bubble-agitated reactors (e.g., bubblecolumn reactors), mechanically agitated reactors (e.g., continuousstirred tank reactors), and flow agitated reactors (e.g., jet reactors).In one embodiment of the invention, the liquid-phase oxidation iscarried out in a single bubble column reactor.

As used herein, the term “bubble column reactor” shall denote a reactorfor facilitating chemical reactions in a multi-phase reaction medium,wherein agitation of the reaction medium is provided primarily by theupward movement of gas bubbles through the reaction medium. As usedherein, the term “agitation” shall denote work dissipated into thereaction medium causing fluid flow and/or mixing. As used herein, theterms “majority”, “primarily”, and “predominately” shall mean more than50 percent. As used herein, the term “mechanical agitation” shall denoteagitation of the reaction medium caused by physical movement of a rigidor flexible element(s) against or within the reaction medium. Forexample, mechanical agitation can be provided by rotation, oscillation,and/or vibration of internal stirrers, paddles, vibrators, or acousticaldiaphragms located in the reaction medium. As used herein, the term“flow agitation” shall denote agitation of the reaction medium caused byhigh velocity injection and/or recirculation of one or more fluids inthe reaction medium. For example, flow agitation can be provided bynozzles, ejectors, and/or eductors.

In a preferred embodiment of the present invention, less than about 40percent of the agitation of the reaction medium in the bubble columnreactor during oxidation is provided by mechanical and/or flowagitation, more preferably less than about 20 percent of the agitationis provided by mechanical and/or flow agitation, and most preferablyless than 5 percent of the agitation is provided by mechanical and/orflow agitation. Preferably, the amount of mechanical and/or flowagitation imparted to the multi-phase reaction medium during oxidationis less than about 3 kilowatts per cubic meter of the reaction medium,more preferably less than about 2 kilowatts per cubic meter, and mostpreferably less than 1 kilowatt per cubic meter.

Referring now to FIG. 1, a preferred bubble column reactor 20 isillustrated as comprising a vessel shell 22 having of a reaction section24 and a disengagement section 26. Reaction section 24 defines aninternal reaction zone 28, while disengagement section 26 defines aninternal disengagement zone 30. A predominately liquid-phase feed streamis introduced into reaction zone 28 via feed inlets 32 a,b,c,d. Apredominately gas-phase oxidant stream is introduced into reaction zone28 via an oxidant sparger 34 located in the lower portion of reactionzone 28. The liquid-phase feed stream and gas-phase oxidant streamcooperatively form a multi-phase reaction medium 36 within reaction zone28. Multi-phase reaction medium 36 comprises a liquid phase and a gasphase. More preferably, multiphase reaction medium 36 comprises athree-phase medium having solid-phase, liquid-phase, and gas-phasecomponents. The solid-phase component of the reaction medium 36preferably precipitates within reaction zone 28 as a result of theoxidation reaction carried out in the liquid phase of reaction medium36. Bubble column reactor 20 includes a slurry outlet 38 located nearthe bottom of reaction zone 28 and a gas outlet 40 located near the topof disengagement zone 30. A slurry effluent comprising liquid-phase andsolid-phase components of reaction medium 36 is withdrawn from reactionzone 28 via slurry outlet 38, while a predominantly gaseous effluent iswithdrawn from disengagement zone 30 via gas outlet 40.

The liquid-phase feed stream introduced into bubble column reactor 20via feed inlets 32 a,b,c,d preferably comprises an oxidizable compound,a solvent, and a catalyst system.

The oxidizable compound present in the liquid-phase feed streampreferably comprises at least one hydrocarbyl group. More preferably,the oxidizable compound is an aromatic compound. Still more preferably,the oxidizable compound is an aromatic compound with at least oneattached hydrocarbyl group or at least one attached substitutedhydrocarbyl group or at least one attached heteroatom or at least oneattached carboxylic acid function (—COOH). Even more preferably, theoxidizable compound is an aromatic compound with at least one attachedhydrocarbyl group or at least one attached substituted hydrocarbyl groupwith each attached group comprising from 1 to 5 carbon atoms. Yet stillmore preferably, the oxidizable compound is an aromatic compound havingexactly two attached groups with each attached group comprising exactlyone carbon atom and consisting of methyl groups and/or substitutedmethyl groups and/or at most one carboxylic acid group. Even still morepreferably, the oxidizable compound is para-xylene, meta-xylene,para-tolualdehyde, meta-tolualdehyde, para-toluic acid, meta-toluicacid, and/or acetaldehyde. Most preferably, the oxidizable compound ispara-xylene.

A “hydrocarbyl group”, as defined herein, is at least one carbon atomthat is bonded only to hydrogen atoms or to other carbon atoms. A“substituted hydrocarbyl group”, as defined herein, is at least onecarbon atom bonded to at least one heteroatom and to at least onehydrogen atom. “Heteroatoms”, as defined herein, are all atoms otherthan carbon and hydrogen atoms. Aromatic compounds, as defined herein,comprise an aromatic ring, preferably having at least 6 carbon atoms,even more preferably having only carbon atoms as part of the ring.Suitable examples of such aromatic rings include, but are not limitedto, benzene, biphenyl, terphenyl, naphthalene, and other carbon-basedfused aromatic rings.

Suitable examples of the oxidizable compound include aliphatichydrocarbons (e.g., alkanes, branched alkanes, cyclic alkanes, aliphaticalkenes, branched alkenes, and cyclic alkenes); aliphatic aldehydes(e.g., acetaldehyde, propionaldehyde, isobutyraldehyde, andn-butyraldehyde); aliphatic alcohols (e.g., ethanol, isopropanol,n-propanol, n-butanol, and isobutanol); aliphatic ketones (e.g.,dimethyl ketone, ethyl methyl ketone, diethyl ketone, and isopropylmethyl ketone); aliphatic esters (e.g., methyl formate, methyl acetate,ethyl acetate); aliphatic peroxides, peracids, and hydroperoxides (e.g.,t-butyl hydroperoxide, peracetic acid, and di-t-butyl hydroperoxide);aliphatic compounds with groups that are combinations of the abovealiphatic species plus other heteroatoms (e.g., aliphatic compoundscomprising one or more molecular segments of hydrocarbons, aldehydes,alcohols, ketones, esters, peroxides, peracids, and/or hydroperoxides incombination with sodium, bromine, cobalt, manganese, and zirconium);various benzene rings, naphthalene rings, biphenyls, terphenyls, andother aromatic groups with one or more attached hydrocarbyl groups(e.g., toluene, ethylbenzene, isopropylbenzene, n-propylbenzene,neopentylbenzene, para-xylene, meta-xylene, ortho-xylene, all isomers oftrimethylbenzenes, all isomers of tetramethylbenzenes,pentamethylbenzene, hexamethylbenzene, all isomers ofethyl-methylbenzenes, all isomers of diethylbenzenes, all isomers ofethyl-dimethylbenzenes, all isomers of dimethylnaphthalenes, all isomersof ethyl-methylnaphthalenes, all isomers of diethylnaphthalenes allisomers of dimethylbiphenyls, all isomers of ethyl-methylbiphenyls, andall isomers of diethylbiphenyls, stilbene and with one or more attachedhydrocarbyl groups, fluorene and with one or more attached hydrocarbylgroups, anthracene and with one or more attached hydrocarbyl groups, anddiphenylethane and with one or more attached hydrocarbyl groups);various benzene rings, naphthalene rings, biphenyls, terphenyls, andother aromatic groups with one or more attached hydrocarbyl groupsand/or one or more attached heteroatoms, which may connect to otheratoms or groups of atoms (e.g., phenol, all isomers of methylphenols,all isomers of dimethylphenols, all isomers of naphthols, benzyl methylether, all isomers of bromophenols, bromobenzene, all isomers ofbromotoluenes including alpha-bromotoluene, dibromobenzene, cobaltnaphthenate, and all isomers of bromobiphenyls); various benzene rings,naphthalene rings, biphenyls, terphenyls, and other aromatic groups withone or more attached hydrocarbyl groups and/or one or more attachedheteroatoms and/or one or more attached substituted hydrocarbyl groups(e.g., benzaldehyde, all isomers of bromobenzaldehydes, all isomers ofbrominated tolualdehydes including all isomers ofalpha-bromotolualdehydes, all isomers of hydroxybenzaldehydes, allisomers of bromo-hydroxybenzaldehydes, all isomers of benzenedicarboxaldehydes, all isomers of benzene tricarboxaldehydes,para-tolualdehyde, meta-tolualdehyde, ortho-tolualdehyde, all isomers oftoluene dicarboxaldehydes, all isomers of toluene tricarboxaldehydes,all isomers of toluene tetracarboxaldehydes, all isomers ofdimethylbenzene dicarboxaldehydes, all isomers of dimethylbenzenetricarboxaldehydes, all isomers of dimethylbenzene tetracarboxaldehydes,all isomers of trimethylbenzene tricarboxaldehydes, all isomers ofethyltolualdehydes, all isomers of trimethylbenzene dicarboxaldehydes,tetramethylbenzene dicarboxaldehyde, hydroxymethyl-benzene, all isomersof hydroxymethyl-toluenes, all isomers of hydroxymethyl-bromotoluenes,all isomers of hydroxymethyl-tolualdehydes, all isomers ofhydroxymethyl-bromotolualdehydes, benzyl hydroperoxide, benzoylhydroperoxide, all isomers of tolyl methyl-hydroperoxides, and allisomers of methylphenol methyl-hydroperoxides); various benzene rings,naphthalenes rings, biphenyls, terphenyls, and other aromatic groupswith one or more attached selected groups, selected groups meaninghydrocarbyl groups and/or attached heteroatoms and/or substitutedhydrocarbyl groups and/or carboxylic acid groups and/or peroxy acidgroups (e.g., benzoic acid, para-toluic acid, meta-toluic acid,ortho-toluic acid, all isomers of ethylbenzoic acids, all isomers ofpropylbenzoic acids, all isomers of butylbenzoic acids, all isomers ofpentylbenzoic acids, all isomers of dimethylbenzoic acids, all isomersof ethylmethylbenzoic acids, all isomers of trimethylbenzoic acids, allisomers of tetramethylbenzoic acids, pentamethylbenzoic acid, allisomers of diethylbenzoic acids, all isomers of benzene dicarboxylicacids, all isomers of benzene tricarboxylic acids, all isomers ofmethylbenzene dicarboxylic acids, all isomers of dimethylbenzenedicarboxylic acids, all isomers of methylbenzene tricarboxylic acids,all isomers of bromobenzoic acids, all isomers of dibromobenzoic acids,all isomers of bromotoluic acids including alpha-bromotoluic acids,tolyl acetic acid, all isomers of hydroxybenzoic acids, all isomers ofhydroxymethyl-benzoic acids, all isomers of hydroxytoluic acids, allisomers of hydroxymethyl-toluic acids, all isomers ofhydroxymethyl-benzene dicarboxylic acids, all isomers ofhydroxybromobenzoic acids, all isomers of hydroxybromotoluic acids, allisomers of hydroxymethyl-bromobenzoic acids, all isomers of carboxybenzaldehydes, all isomers of dicarboxy benzaldehydes, perbenzoic acid,all isomers of hydroperoxymethyl-benzoic acids, all isomers ofhydroperoxymethyl-hydroxybenzoic acids, all isomers ofhydroperoxycarbonyl-benzoic acids, all isomers ofhydroperoxycarbonyl-toluenes, all isomers of methylbiphenyl carboxylicacids, all isomers of dimethylbiphenyl carboxylic acids, all isomers ofmethylbiphenyl dicarboxylic acids, all isomers of biphenyl tricarboxylicacids, all isomers of stilbene with one or more attached selectedgroups, all isomers of fluorenone with one or more attached selectedgroups, all isomers of naphthalene with one or more attached selectedgroups, benzil, all isomers of benzil with one or more attached selectedgroups, benzophenone, all isomers of benzophenone with one or moreattached selected groups, anthraquinone, all isomers of anthraquinonewith one or more attached selected groups, all isomers of diphenylethanewith one or more attached selected groups, benzocoumarin, and allisomers of benzocoumarin with one or more attached selected groups).

If the oxidizable compound present in the liquid-phase feed stream is anormally-solid compound (i.e., is a solid at standard temperature andpressure), it is preferred for the oxidizable compound to besubstantially dissolved in the solvent when introduced into reactionzone 28. It is preferred for the boiling point of the oxidizablecompound at atmospheric pressure to be at least about 50° C. Morepreferably, the boiling point of the oxidizable compound is in the rangeof from about 80 to about 400° C., and most preferably in the range offrom 125 to 155° C. The amount of oxidizable compound present in theliquid-phase feed is preferably in the range of from about 2 to about 40weight percent, more preferably in the range of from about 4 to about 20weight percent, and most preferably in the range of from 6 to 15 weightpercent.

It is now noted that the oxidizable compound present in the liquid-phasefeed may comprise a combination of two or more different oxidizablechemicals. These two or more different chemical materials can be fedcommingled in the liquid-phase feed stream or may be fed separately inmultiple feed streams. For example, an oxidizable compound comprisingpara-xylene, meta-xylene, para-tolualdehyde, para-toluic acid, andacetaldehyde may be fed to the reactor via a single inlet or multipleseparate inlets.

The solvent present in the liquid-phase feed stream preferably comprisesan acid component and a water component. The solvent is preferablypresent in the liquid-phase feed stream at a concentration in the rangeof from about 60 to about 98 weight percent, more preferably in therange of from about 80 to about 96 weight percent, and most preferablyin the range of from 85 to 94 weight percent. The acid component of thesolvent is preferably primarily an organic low molecular weightmonocarboxylic acid having 1-6 carbon atoms, more preferably 2 carbonatoms. Most preferably, the acid component of the solvent is primarilyacetic acid. Preferably, the acid component makes up at least about 75weight percent of the solvent, more preferably at least about 80 weightpercent of the solvent, and most preferably 85 to 98 weight percent ofthe solvent, with the balance being primarily water. The solventintroduced into bubble column reactor 20 can include small quantities ofimpurities such as, for example, para-tolualdehyde, terephthaldehyde,4-carboxybenzaldehyde (4-CBA), benzoic acid, para-toluic acid,para-toluic aldehyde, alpha-bromo-para-toluic acid, isophthalic acid,phthalic acid, trimellitic acid, polyaromatics, and/or suspendedparticulate. It is preferred that the total amount of impurities in thesolvent introduced into bubble column reactor 20 is less than about 3weight percent.

The catalyst system present in the liquid-phase feed stream ispreferably a homogeneous, liquid-phase catalyst system capable ofpromoting oxidation (including partial oxidation) of the oxidizablecompound. More preferably, the catalyst system comprises at least onemultivalent transition metal. Still more preferably, the multivalenttransition metal comprises cobalt. Even more preferably, the catalystsystem comprises cobalt and bromine. Most preferably, the catalystsystem comprises cobalt, bromine, and manganese.

When cobalt is present in the catalyst system, it is preferred for theamount of cobalt present in the liquid-phase feed stream to be such thatthe concentration of cobalt in the liquid phase of reaction medium 36 ismaintained in the range of from about 300 to about 6,000 parts permillion by weight (ppmw), more preferably in the range of from about 700to about 4,200 ppmw, and most preferably in the range of from 1,200 to3,000 ppmw. When bromine is present in the catalyst system, it ispreferred for the amount of bromine present in the liquid-phase feedstream to be such that the concentration of bromine in the liquid phaseof reaction medium 36 is maintained in the range of from about 300 toabout 5,000 ppmw, more preferably in the range of from about 600 toabout 4,000 ppmw, and most preferably in the range of from 900 to 3,000ppmw. When manganese is present in the catalyst system, it is preferredfor the amount of manganese present in the liquid-phase feed stream tobe such that the concentration of manganese in the liquid phase ofreaction medium 36 is maintained in the range of from about 20 to about1,000 ppmw, more preferably in the range of from about 40 to about 500ppmw, most preferably in the range of from 50 to 200 ppmw.

The concentrations of the cobalt, bromine, and/or manganese in theliquid phase of reaction medium 36, provided above, are expressed on atime-averaged and volume-averaged basis. As used herein, the term“time-averaged” shall denote an average of at least 10 measurementstaken equally over a continuous period of at least 100 seconds. As usedherein, the term “volume-averaged” shall denote an average of at least10 measurements taken at uniform 3-dimensional spacing throughout acertain volume.

The weight ratio of cobalt to bromine (Co:Br) in the catalyst systemintroduced into reaction zone 28 is preferably in the range of fromabout 0.25:1 to about 4:1, more preferably in the range of from about0.5:1 to about 3:1, and most preferably in the range of from 0.75:1 to2:1. The weight ratio of cobalt to manganese (Co:Mn) in the catalystsystem introduced into reaction zone 28 is preferably in the range offrom about 0.3:1 to about 40:1, more preferably in the range of fromabout 5:1 to about 30:1, and most preferably in the range of from 10:1to 25:1.

The liquid-phase feed stream introduced into bubble column reactor 20can include small quantities of impurities such as, for example,toluene, ethylbenzene, para-tolualdehyde, terephthaldehyde,4-carboxybenzaldehyde (4-CBA), benzoic acid, para-toluic acid,para-toluic aldehyde, alpha bromo para-toluic acid, isophthalic acid,phthalic acid, trimellitic acid, polyaromatics, and/or suspendedparticulate. When bubble column reactor 20 is employed for theproduction of terephthalic acid, meta-xylene and ortho-xylene are alsoconsidered impurities. It is preferred that the total amount ofimpurities in the liquid-phase feed stream introduced into bubble columnreactor 20 is less than about 3 weight percent.

Although FIG. 1 illustrates an embodiment where the oxidizable compound,the solvent, and the catalyst system are mixed together and introducedinto bubble column reactor 20 as a single feed stream, in an alternativeembodiment of the present invention, the oxidizable compound, thesolvent, and the catalyst can be separately introduced into bubblecolumn reactor 20. For example, it is possible to feed a purepara-xylene stream into bubble column reactor 20 via an inlet separatefrom the solvent and catalyst inlet(s).

The predominately gas-phase oxidant stream introduced into bubble columnreactor 20 via oxidant sparger 34 comprises molecular oxygen (O₂).Preferably, the oxidant stream comprises in the range of from about 5 toabout 40 mole percent molecular oxygen, more preferably in the range offrom about 15 to about 30 mole percent molecular oxygen, and mostpreferably in the range of from 18 to 24 mole percent molecular oxygen.It is preferred for the balance of the oxidant stream to be comprisedprimarily of a gas or gasses, such as nitrogen, that are inert tooxidation. More preferably, the oxidant stream consists essentially ofmolecular oxygen and nitrogen. Most preferably, the oxidant stream isdry air that comprises about 21 mole percent molecular oxygen and about78 to about 81 mole percent nitrogen. In an alternative embodiment ofthe present invention, the oxidant stream can comprise substantiallypure oxygen.

Referring again to FIG. 1, bubble column reactor 20 is preferablyequipped with a reflux distributor 42 positioned above an upper surface44 of reaction medium 36. Reflux distributor 42 is operable to introducedroplets of a predominately liquid-phase reflux stream intodisengagement zone 30 by any means of droplet formation known in theart. More preferably, reflux distributor 42 produces a spray of dropletsdirected downwardly towards upper surface 44 of reaction medium 36.Preferably, this downward spray of droplets affects (i.e., engages andinfluences) at least about 50 percent of the maximum horizontalcross-sectional area of disengagement zone 30. More preferably, thespray of droplets affects at least about 75 percent of the maximumhorizontal cross-sectional area of disengagement zone 30. Mostpreferably, the spray of droplets affects at least 90 percent of themaximum horizontal cross-sectional area of disengagement zone 30. Thisdownward liquid reflux spray can help prevent foaming at or above uppersurface 44 of reaction medium 36 and can also aid in the disengagementof any liquid or slurry droplets entrained in the upwardly moving gasthat flows towards gas outlet 40. Further, the liquid reflux may serveto reduce the amount of particulates and potentially precipitatingcompounds (e.g., dissolved benzoic acid, para-toluic acid, 4-CBA,terephthalic acid, and catalyst metal salts) exiting in the gaseouseffluent withdrawn from disengagement zone 30 via gas outlet 40. Inaddition, the introduction of reflux droplets into disengagement zone 30can, by a distillation action, be used to adjust the composition of thegaseous effluent withdrawn via gas outlet 40.

The liquid reflux stream introduced into bubble column reactor 20 viareflux distributor 42 preferably has about the same composition as thesolvent component of the liquid-phase feed stream introduced into bubblecolumn reactor 20 via feed inlets 32 a,b,c,d. Thus, it is preferred forthe liquid reflux stream to comprise an acid component and water. Theacid component of the reflux stream is preferably a low molecular weightorganic monocarboxylic acid having 1-6 carbon atoms, more preferably 2carbon atoms. Most preferably, the acid component of the reflux streamis acetic acid. Preferably, the acid component makes up at least about75 weight percent of the reflux stream, more preferably at least about80 weight percent of the reflux stream, and most preferably 85 to 98weight percent of the reflux stream, with the balance being water.Because the reflux stream typically has substantially the samecomposition as the solvent in the liquid-phase feed stream, when thisdescription refers to the “total solvent” introduced into the reactor,such “total solvent” shall include both the reflux stream and thesolvent portion of the feed stream.

During liquid-phase oxidation in bubble column reactor 20, it ispreferred for the feed, oxidant, and reflux streams to be substantiallycontinuously introduced into reaction zone 28, while the gas and slurryeffluent streams are substantially continuously withdrawn from reactionzone 28. As used herein, the term “substantially continuously” shallmean for a period of at least 10 hours interrupted by less than 10minutes. During oxidation, it is preferred for the oxidizable compound(e.g., para-xylene) to be substantially continuously introduced intoreaction zone 28 at a rate of at least about 8,000 kilograms per hour,more preferably at a rate in the range of from about 13,000 to about80,000 kilograms per hour, still more preferably in the range of fromabout 18,000 to about 50,000 kilograms per hour, and most preferably inthe range of from 22,000 to 30,000 kilograms per hour. Although it isgenerally preferred for the flow rates of the incoming feed, oxidant,and reflux streams to be substantially steady, it is now noted that oneembodiment of the presenting invention contemplates pulsing the incomingfeed, oxidant, and/or reflux stream in order to improve mixing and masstransfer. When the incoming feed, oxidant, and/or reflux stream areintroduced in a pulsed fashion, it is preferred for their flow rates tovary within about 0 to about 500 percent of the steady-state flow ratesrecited herein, more preferably within about 30 to about 200 percent ofthe steady-state flow rates recited herein, and most preferably within80 to 120 percent of the steady-state flow rates recited herein.

The average space-time rate of reaction (STR) in bubble column oxidationreactor 20 is defined as the mass of the oxidizable compound fed perunit volume of reaction medium 36 per unit time (e.g., kilograms ofpara-xylene fed per cubic meter per hour). In conventional usage, theamount of oxidizable compound not converted to product would typicallybe subtracted from the amount of oxidizable compound in the feed streambefore calculating the STR. However, conversions and yields aretypically high for many of the oxidizable compounds preferred herein(e.g., para-xylene), and it is convenient to define the term herein asstated above. For reasons of capital cost and operating inventory, amongothers, it is generally preferred that the reaction be conducted with ahigh STR. However, conducting the reaction at increasingly higher STRmay affect the quality or yield of the partial oxidation. Bubble columnreactor 20 is particularly useful when the STR of the oxidizablecompound (e.g., para-xylene) is in the range of from about 25 kilogramsper cubic meter per hour to about 400 kilograms per cubic meter perhour, more preferably in the range of from about 30 kilograms per cubicmeter per hour to about 250 kilograms per cubic meter per hour, stillmore preferably from about 35 kilograms per cubic meter per hour toabout 150 kilograms per cubic meter per hour, and most preferably in therange of from 40 kilograms per cubic meter per hour to 100 kilograms percubic meter per hour.

The oxygen-STR in bubble column oxidation reactor 20 is defined as theweight of molecular oxygen consumed per unit volume of reaction medium36 per unit time (e.g., kilograms of molecular oxygen consumed per cubicmeter per hour). For reasons of capital cost and oxidative consumptionof solvent, among others, it is generally preferred that the reaction beconducted with a high oxygen-STR. However, conducting the reaction atincreasingly higher oxygen-STR eventually reduces the quality or yieldof the partial oxidation. Without being bound by theory, it appears thatthis possibly relates to the transfer rate of molecular oxygen from thegas phase into the liquid at the interfacial surface area and thenceinto the bulk liquid. Too high an oxygen-STR possibly leads to too low adissolved oxygen content in the bulk liquid phase of the reactionmedium.

The global-average-oxygen-STR is defined herein as the weight of alloxygen consumed in the entire volume of reaction medium 36 per unit time(e.g., kilograms of molecular oxygen consumed per cubic meter per hour).Bubble column reactor 20 is particularly useful when theglobal-average-oxygen-STR is in the range of from about 25 kilograms percubic meter per hour to about 400 kilograms per cubic meter per hour,more preferably in the range of from about 30 kilograms per cubic meterper hour to about 250 kilograms per cubic meter per hour, still morepreferably from about 35 kilograms per cubic meter per hour to about 150kilograms per cubic meter per hour, and most preferably in the range offrom 40 kilograms per cubic meter per hour to 100 kilograms per cubicmeter per hour.

During oxidation in bubble column reactor 20, it is preferred for theratio of the mass flow rate of the total solvent (from both the feed andreflux streams) to the mass flow rate of the oxidizable compoundentering reaction zone 28 to be maintained in the range of from about2:1 to about 50:1, more preferably in the range of from about 5:1 toabout 40:1, and most preferably in the range of from 7.5:1 to 25:1.Preferably, the ratio of the mass flow rate of solvent introduced aspart of the feed stream to the mass flow rate of solvent introduced aspart of the reflux stream is maintained in the range of from about 0.5:1to no reflux stream flow whatsoever, more preferably in the range offrom about 0.5:1 to about 4:1, still more preferably in the range offrom about 1:1 to about 2:1, and most preferably in the range of from1.25:1 to 1.5:1.

During liquid-phase oxidation in bubble column reactor 20, it ispreferred for the oxidant stream to be introduced into bubble columnreactor 20 in an amount that provides molecular oxygen somewhatexceeding the stoichiometric oxygen demand. The amount of excessmolecular oxygen required for best results with a particular oxidizablecompound affects the overall economics of the liquid-phase oxidation.During liquid-phase oxidation in bubble column reactor 20, it ispreferred that the ratio of the mass flow rate of the oxidant stream tothe mass flow rate of the oxidizable organic compound (e.g.,para-xylene) entering reactor 20 is maintained in the range of fromabout 0.5:1 to about 20:1, more preferably in the range of from about1:1 to about 10:1, and most preferably in the range of from 2:1 to 6:1.

Referring again to FIG. 1, the feed, oxidant, and reflux streamsintroduced into bubble column reactor 20 cooperatively form at least aportion of multi-phase reaction medium 36. Reaction medium 36 ispreferably a three-phase medium comprising a solid phase, a liquidphase, and a gas phase. As mentioned above, oxidation of the oxidizablecompound (e.g., para-xylene) takes place predominately in the liquidphase of reaction medium 36. Thus, the liquid phase of reaction medium36 comprises dissolved oxygen and the oxidizable compound. Theexothermic nature of the oxidation reaction that takes place in bubblecolumn reactor 20 causes a portion of the solvent (e.g., acetic acid andwater) introduced via feed inlets 32 a,b,c,d to boil/vaporize. Thus, thegas phase of reaction medium 36 in reactor 20 is formed primarily ofvaporized solvent and an undissolved, unreacted portion of the oxidantstream. Certain prior art oxidation reactors employ heat exchangetubes/fins to heat or cool the reaction medium. However, such heatexchange structures may be undesirable in the inventive reactor andprocess described herein. Thus, it is preferred for bubble columnreactor 20 to include substantially no surfaces that contact reactionmedium 36 and exhibit a time-averaged heat flux greater than 30,000watts per meter squared.

The concentration of dissolved oxygen in the liquid phase of reactionmedium 36 is a dynamic balance between the rate of mass transfer fromthe gas phase and the rate of reactive consumption within the liquidphase (i.e. it is not set simply by the partial pressure of molecularoxygen in the supplying gas phase, though this is one factor in thesupply rate of dissolved oxygen and it does affect the limiting upperconcentration of dissolved oxygen). The amount of dissolved oxygenvaries locally, being higher near bubble interfaces. Globally, theamount of dissolved oxygen depends on the balance of supply and demandfactors in different regions of reaction medium 36. Temporally, theamount of dissolved oxygen depends on the uniformity of gas and liquidmixing relative to chemical consumption rates. In designing to matchappropriately the supply of and demand for dissolved oxygen in theliquid phase of reaction medium 36, it is preferred for thetime-averaged and volume-averaged oxygen concentration in the liquidphase of reaction medium 36 to be maintained above about 1 ppm molar,more preferably in the range from about 4 to about 1,000 ppm molar,still more preferably in the range from about 8 to about 500 ppm molar,and most preferably in the range from 12 to 120 ppm molar.

The liquid-phase oxidation reaction carried out in bubble column reactor20 is preferably a precipitating reaction that generates solids. Morepreferably, the liquid-phase oxidation carried out in bubble columnreactor 20 causes at least about 10 weight percent of the oxidizablecompound (e.g., para-xylene) introduced into reaction zone 28 to form asolid compound (e.g., crude terephthalic acid particles) in reactionmedium 36. Still more preferably, the liquid-phase oxidation causes atleast about 50 weight percent of the oxidizable compound to form a solidcompound in reaction medium 36. Most preferably, the liquid-phaseoxidation causes at least 90 weight percent of the oxidizable compoundto form a solid compound in reaction medium 36. It is preferred for thetotal amount of solids in reaction medium 36 to be greater than about 3percent by weight on a time-averaged and volume-averaged basis. Morepreferably, the total amount of solids in reaction medium 36 ismaintained in the range of from about 5 to about 40 weight percent,still more preferably in the range of from about 10 to about 35 weightpercent, and most preferably in the range of from 15 to 30 weightpercent. It is preferred for a substantial portion of the oxidationproduct (e.g., terephthalic acid) produced in bubble column reactor 20to be present in reaction medium 36 as solids, as opposed to remainingdissolved in the liquid phase of reaction medium 36. The amount of thesolid phase oxidation product present in reaction medium 36 ispreferably at least about 25 percent by weight of the total oxidationproduct (solid and liquid phase) in reaction medium 36, more preferablyat least about 75 percent by weight of the total oxidation product inreaction medium 36, and most preferably at least 95 percent by weight ofthe total oxidation product in reaction medium 36. The numerical rangesprovided above for the amount of solids in reaction medium 36 apply tosubstantially steady-state operation of bubble column 20 over asubstantially continuous period of time, not to start-up, shut-down, orsub-optimal operation of bubble column reactor 20. The amount of solidsin reaction medium 36 is determined by a gravimetric method. In thisgravimetric method, a representative portion of slurry is withdrawn fromthe reaction medium and weighed. At conditions that effectively maintainthe overall solid-liquid partitioning present within the reactionmedium, free liquid is removed from the solids portion by sedimentationor filtration, effectively without loss of precipitated solids and withless than about 10 percent of the initial liquid mass remaining with theportion of solids. The remaining liquid on the solids is evaporated todryness, effectively without sublimation of solids. The remainingportion of solids is weighed. The ratio of the weight of the portion ofsolids to the weight of the original portion of slurry is the fractionof solids, typically expressed as a percentage.

The precipitating reaction carried out in bubble column reactor 20 cancause fouling (i.e., solids build-up) on the surface of certain rigidstructures that contact reaction medium 36. Thus, in one embodiment ofthe present invention, it is preferred for bubble column reactor 20 toinclude substantially no internal heat exchange, stirring, or bafflingstructures in reaction zone 28 because such structures would be prone tofouling. If internal structures are present in reaction zone 28, it isdesirable to avoid internal structures having outer surfaces thatinclude a significant amount of upwardly facing planar surface areabecause such upwardly facing planar surfaces would be highly prone tofouling. Thus, if any internal structures are present in reaction zone28, it is preferred for less than about 20 percent of the total upwardlyfacing exposed outer surface area of such internal structures to beformed by substantially planar surfaces inclined less than about 15degrees from horizontal.

Referring again to FIG. 1, the physical configuration of bubble columnreactor 20 helps provide for optimized oxidation of the oxidizablecompound (e.g., para-xylene) with minimal impurity generation. It ispreferred for elongated reaction section 24 of vessel shell 22 toinclude a substantially cylindrical main body 46 and a lower head 48.The upper end of reaction zone 28 is defined by a horizontal plane 50extending across the top of cylindrical main body 46. A lower end 52 ofreaction zone 28 is defined by the lowest internal surface of lower head48. Typically, lower end 52 of reaction zone 28 is located proximate theopening for slurry outlet 38. Thus, elongated reaction zone 28 definedwithin bubble column reactor 20 has a maximum length “L” measured fromthe top end 50 to the bottom end 52 of reaction zone 28 along the axisof elongation of cylindrical main body 46. The length “L” of reactionzone 28 is preferably in the range of from about 10 to about 100 meters,more preferably in the range of from about 20 to about 75 meters, andmost preferably in the range of from 25 to 50 meters. Reaction zone 28has a maximum diameter (width) “D” that is typically equal to themaximum internal diameter of cylindrical main body 46. The maximumdiameter “D” of reaction zone 28 is preferably in the range of fromabout 1 to about 12 meters, more preferably in the range of from about 2to about 10 meters, still more preferably in the range of from about 3.1to about 9 meters, and most preferably in the range of from 4 to 8meters. In a preferred embodiment of the present invention, reactionzone 28 has a length-to-diameter “L:D” ratio in the range of from about6:1 to about 30:1. Still more preferably, reaction zone 28 has an L:Dratio in the range of from about 8:1 to about 20:1. Most preferably,reaction zone 28 has an L:D ratio in the range of from 9:1 to 15:1.

As discussed above, reaction zone 28 of bubble column reactor 20receives multi-phase reaction medium 36. Reaction medium 36 has a bottomend coincident with lower end 52 of reaction zone 28 and a top endlocated at upper surface 44. Upper surface 44 of reaction medium 36 isdefined along a horizontal plane that cuts through reaction zone 28 at avertical location where the contents of reaction zone 28 transitionsfrom a gas-phase-continuous state to a liquid-phase-continuous state.Upper surface 44 is preferably positioned at the vertical location wherethe local time-averaged gas hold-up of a thin horizontal slice of thecontents of reaction zone 28 is 0.9.

Reaction medium 36 has a maximum height “H” measured between its upperand lower ends. The maximum width “W” of reaction medium 36 is typicallyequal to the maximum diameter “D” of cylindrical main body 46. Duringliquid-phase oxidation in bubble column reactor 20, it is preferred thatH is maintained at about 60 to about 120 percent of L, more preferablyabout 80 to about 110 percent of L, and most preferably 85 to 100percent of L. In a preferred embodiment of the present invention,reaction medium 36 has a height-to-width “H:W” ratio greater than about3:1. More preferably, reaction medium 36 has an H:W ratio in the rangeof from about 7:1 to about 25:1. Still more preferably, reaction medium36 has an H:W ratio in the range of from about 8:1 to about 20:1. Mostpreferably, reaction medium 36 has an H:W ratio in the range of from 9:1to 15:1. In one embodiment of the invention, L=H and D=W so that variousdimensions or ratios provide herein for L and D also apply to H and W,and vice-versa.

The relatively high L:D and H:W ratios provided in accordance with anembodiment of the invention can contribute to several importantadvantages of the inventive system. As discussed in further detailbelow, it has been discovered that higher L:D and H:W ratios, as well ascertain other features discussed below, can promote beneficial verticalgradients in the concentrations of molecular oxygen and/or theoxidizable compound (e.g., para-xylene) in reaction medium 36. Contraryto conventional wisdom, which would favor a well-mixed reaction mediumwith relatively uniform concentrations throughout, it has beendiscovered that the vertical staging of the oxygen and/or the oxidizablecompound concentrations facilitates a more effective and economicaloxidation reaction. Minimizing the oxygen and oxidizable compoundconcentrations near the top of reaction medium 36 can help avoid loss ofunreacted oxygen and unreacted oxidizable compound through upper gasoutlet 40. However, if the concentrations of oxidizable compound andunreacted oxygen are low throughout reaction medium 36, then the rateand/or selectivity of oxidation are reduced. Thus, it is preferred forthe concentrations of molecular oxygen and/or the oxidizable compound tobe significantly higher near the bottom of reaction medium 36 than nearthe top of reaction medium 36.

In addition, high L:D and H:W ratios cause the pressure at the bottom ofreaction medium 36 to be substantially greater than the pressure at thetop of reaction medium 36. This vertical pressure gradient is a resultof the height and density of reaction medium 36. One advantage of thisvertical pressure gradient is that the elevated pressure at the bottomof the vessel drives more oxygen solubility and mass transfer than wouldotherwise be achievable at comparable temperatures and overheadpressures in shallow reactors. Thus, the oxidation reaction can becarried out at lower temperatures than would be required in a shallowervessel. When bubble column reactor 20 is used for the partial oxidationof para-xylene to crude terephthalic acid (CTA), the ability to operateat lower reaction temperatures with the same or better oxygen masstransfer rates has a number of advantages. For example, low temperatureoxidation of para-xylene reduces the amount of solvent burned during thereaction. As discussed in further detail below, low temperatureoxidation also favors the formation of small, high surface area, looselybound, easily dissolved CTA particles, which can be subjected to moreeconomical purification techniques than the large, low surface area,dense CTA particles produced by conventional high temperature oxidationprocesses.

During oxidation in reactor 20, it is preferred for the time-averagedand volume-averaged temperature of reaction medium 36 to be maintainedin the range of from about 125 to about 200° C., more preferably in therange of from about 140 to about 180° C., and most preferably in therange of from 150 to 170° C. The overhead pressure above reaction medium36 is preferably maintained in the range of from about 1 to about 20 bargauge (barg), more preferably in the range of from about 2 to about 12barg, and most preferably in the range of from 4 to 8 barg. Preferably,the pressure difference between the top of reaction medium 36 and thebottom of reaction medium 36 is in the range of from about 0.4 to about5 bar, more preferably the pressure difference is in the range of fromabout 0.7 to about 3 bars, and most preferably the pressure differenceis 1 to 2 bar. Although it is generally preferred for the overheadpressure above reaction medium 36 to be maintained at a relativelyconstant value, one embodiment of the present invention contemplatespulsing the overhead pressure to facilitate improved mixing and/or masstransfer in reaction medium 36. When the overhead pressure is pulsed, itis preferred for the pulsed pressures to range between about 60 to about140 percent of the steady-state overhead pressure recited herein, morepreferably between about 85 and about 115 percent of the steady-stateoverhead pressure recited herein, and most preferably between 95 and 105percent of the steady-state overhead pressure recited herein.

A further advantage of the high L:D ratio of reaction zone 28 is that itcan contribute to an increase in the average superficial velocity ofreaction medium 36. The term “superficial velocity” and “superficial gasvelocity”, as used herein with reference to reaction medium 36, shalldenote the volumetric flow rate of the gas phase of reaction medium 36at an elevation in the reactor divided by the horizontal cross-sectionalarea of the reactor at that elevation. The increased superficialvelocity provided by the high L:D ratio of reaction zone 28 can promotelocal mixing and increase the gas hold-up of reaction medium 36. Thetime-averaged superficial velocities of reaction medium 36 atone-quarter height, half height, and/or three-quarter height of reactionmedium 36 are preferably greater than about 0.3 meters per second, morepreferably in the range of from about 0.8 to about 5 meters per second,still more preferably in the range of from about 0.9 to about 4 metersper second, and most preferably in the range of from 1 to 3 meters persecond.

Referring again to FIG. 1, disengagement section 26 of bubble columnreactor 20 is simply a widened portion of vessel shell 22 locatedimmediately above reaction section 24. Disengagement section 26 reducesthe velocity of the upwardly-flowing gas phase in bubble column reactor20 as the gas phase rises above the upper surface 44 of reaction medium36 and approaches gas outlet 40. This reduction in the upward velocityof the gas phase helps facilitate removal of entrained liquids and/orsolids in the upwardly flowing gas phase and thereby reduces undesirableloss of certain components present in the liquid phase of reactionmedium 36.

Disengagement section 26 preferably includes a generally frustoconicaltransition wall 54, a generally cylindrical broad sidewall 56, and anupper head 58. The narrow lower end of transition wall 54 is coupled tothe top of cylindrical main body 46 of reaction section 24. The wideupper end of transition wall 54 is coupled to the bottom of broadsidewall 56. It is preferred for transition wall 54 to extend upwardlyand outwardly from its narrow lower end at an angle in the range of fromabout 10 to about 70 degrees from vertical, more preferably in the rangeof about 15 to about 50 degrees from vertical, and most preferably inthe range of from 15 to 45 degrees from vertical. Broad sidewall 56 hasa maximum diameter “X” that is generally greater than the maximumdiameter “D” of reaction section 24, though when the upper portion ofreaction section 24 has a smaller diameter than the overall maximumdiameter of reaction section 24, then X may actually be smaller than D.In a preferred embodiment of the present invention, the ratio of thediameter of broad sidewall 56 to the maximum diameter of reactionsection 24 “X:D” is in the range of from about 0.8:1 to about 4:1, mostpreferably in the range of from 1.1:1 to 2:1. Upper head 58 is coupledto the top of broad sidewall 56. Upper head 58 is preferably a generallyelliptical head member defining a central opening that permits gas toescape disengagement zone 30 via gas outlet 40. Alternatively, upperhead 58 may be of any shape, including conical. Disengagement zone 30has a maximum height “Y” measured from the top 50 of reaction zone 28 tothe upper most portion of disengagement zone 30. The ratio of the lengthof reaction zone 28 to the height of disengagement zone 30 “L:Y” ispreferably in the range of from about 2:1 to about 24:1, more preferablyin the range of from about 3:1 to about 20:1, and most preferably in therange of from 4:1 to 16:1.

Referring now to FIGS. 1-5, the location and configuration of oxidantsparger 34 will now be discussed in greater detail. FIGS. 2 and 3 showthat oxidant sparger 34 can include a ring member 60, a cross-member 62,and a pair of oxidant entry conduits 64 a,b. Conveniently, these oxidantentry conduits 64 a,b can enter the vessel at an elevation above thering member 60 and then turn downwards as shown in FIGS. 2 and 3.Alternatively, an oxidant entry conduit 64 a,b may enter the vesselbelow the ring member 60 or on about the same horizontal plane as ringmember 60. Each oxidant entry conduit 64 a,b includes a first endcoupled to a respective oxidant inlet 66 a,b formed in the vessel shell22 and a second end fluidly coupled to ring member 60. Ring member 60 ispreferably formed of conduits, more preferably of a plurality ofstraight conduit sections, and most preferably a plurality of straightpipe sections, rigidly coupled to one another to form a tubularpolygonal ring. Preferably, ring member 60 is formed of at least 3straight pipe sections, more preferably 6 to 10 pipe sections, and mostpreferably 8 pipe sections. Accordingly, when ring member 60 is formedof 8 pipe sections, it has a generally octagonal configuration.Cross-member 62 is preferably formed of a substantially straight pipesection that is fluidly coupled to and extends diagonally betweenopposite pipe sections of ring member 60. The pipe section used forcross-member 62 preferably has substantially the same diameter as thepipe sections used to form ring member 60. It is preferred for the pipesections that make up oxidant entry conduits 64 a,b, ring member 60, andcross-member 62 to have a nominal diameter greater than about 0.1 meter,more preferable in the range of from about 0.2 to about 2 meters, andmost preferably in the range of from 0.25 to 1 meters. As perhaps bestillustrated in FIG. 3, ring member 60 and cross-member 62 each present aplurality of upper oxidant openings 68 for discharging the oxidantstream upwardly into reaction zone 28. As perhaps best illustrated inFIG. 4, ring member 60 and/or cross-member 62 can present one or morelower oxidant openings 70 for discharging the oxidant stream downwardlyinto reaction zone 28. Lower oxidant openings 70 can also be used todischarge liquids and/or solids that might intrude within ring member 60and/or cross-member 62. In order to prevent solids from building upinside oxidant sparger 34, a liquid stream can be continuously orperiodically passed through sparger 34 to flush out any accumulatedsolids.

Referring again to FIGS. 1-4, during oxidation in bubble column reactor20, oxidant streams are forced through oxidant inlets 66 a,b and intooxidant entry conduits 64 a,b, respectively. The oxidant streams arethen transported via oxidant entry conduits 64 a,b to ring member 60.Once the oxidant stream has entered ring member 60, the oxidant streamis distributed throughout the internal volumes of ring member 60 andcross-member 62. The oxidant stream is then forced out of oxidantsparger 34 and into reaction zone 28 via upper and lower oxidantopenings 68,70 of ring member 60 and cross-member 62.

The outlets of upper oxidant openings 68 are laterally spaced from oneanother and are positioned at substantially the same elevation inreaction zone 28. Thus, the outlets of upper oxidant openings 68 aregenerally located along a substantially horizontal plane defined by thetop of oxidant sparger 34. The outlets of lower oxidant openings 70 arelaterally spaced from one another and are positioned at substantiallythe same elevation in reaction zone 28. Thus, the outlets of loweroxidant openings 70 are generally located along a substantiallyhorizontal plane defined by the bottom of oxidant sparger 34.

In one embodiment of the present invention, oxidant sparger 34 has atleast about 20 upper oxidant openings 68 formed therein. Morepreferably, oxidant sparger 34 has in the range of from about 40 toabout 800 upper oxidant openings formed therein. Most preferably,oxidant sparger 34 has in the range of from 60 to 400 upper oxidantopenings 68 formed therein. Oxidant sparger 34 preferably has at leastabout 1 lower oxidant opening 70 formed therein. More preferably,oxidant sparger 34 has in the range of from about 2 to about 40 loweroxidant openings 70 formed therein. Most preferably, oxidant sparger 34has in the range of from 8 to 20 lower oxidant openings 70 formedtherein. The ratio of the number of upper oxidant openings 68 to loweroxidant openings 70 in oxidant sparger 34 is preferably in the range offrom about 2:1 to about 100:1, more preferably in the range of fromabout 5:1 to about 25:1, and most preferably in the range of from 8:1 to15:1. The diameters of substantially all upper and lower oxidantopenings 68,70 are preferably substantially the same, so that the ratioof the volumetric flow rate of the oxidant stream out of upper and loweropenings 68,70 is substantially the same as the ratios, given above, forthe relative number of upper and lower oxidant openings 68,70.

FIG. 5 illustrates the direction of oxidant discharge from upper andlower oxidant openings 68,70. With reference to upper oxidant openings68, it is preferred for at least a portion of upper oxidant openings 68to discharge the oxidant stream in at an angle “A” that is skewed fromvertical. It is preferred for the percentage of upper oxidant openings68 that are skewed from vertical by angle “A” to be in the range of fromabout 30 to about 90 percent, more preferably in the range of from about50 to about 80 percent, still more preferably in the range of from 60 to75 percent, and most preferably about 67 percent. The angle “A” ispreferably in the range of from about 5 to about 60 degrees, morepreferably in the range of from about 10 to about 45 degrees, and mostpreferably in the range of from 15 to 30 degrees. As for lower oxidantopenings 70, it is preferred that substantially all of lower oxidantopenings 70 are located near the bottom-most portion of the ring member60 and/or cross-member 62. Thus, any liquids and/or solids that may haveunintentionally entered oxidant sparger 34 can be readily dischargedfrom oxidant sparger 34 via lower oxidant openings 70. Preferably, loweroxidant openings 70 discharge the oxidant stream downwardly at asubstantially vertical angle. For purposes of this description, an upperoxidant opening can be any opening that discharges an oxidant stream ina generally upward direction (i.e., at an angle above horizontal), and alower oxidant opening can be any opening that discharges an oxidantstream in a generally downward direction (i.e., at an angle belowhorizontal).

In many conventional bubble column reactors containing a multi-phasereaction medium, substantially all of the reaction medium located belowthe oxidant sparger (or other mechanism for introducing the oxidantstream into the reaction zone) has a very low gas hold-up value. Asknown in the art, “gas hold-up” is simply the volume fraction of amulti-phase medium that is in the gaseous state. Zones of low gashold-up in a medium can also be referred to as “unaerated” zones. Inmany conventional slurry bubble column reactors, a significant portionof the total volume of the reaction medium is located below the oxidantsparger (or other mechanism for introducing the oxidant stream into thereaction zone). Thus, a significant portion of the reaction mediumpresent at the bottom of conventional bubble column reactors isunaerated.

It has been discovered that minimizing the amount of unaerated zones ina reaction medium subjected to oxidization in a bubble column reactorcan minimize the generation of certain types of undesirable impurities.Unaerated zones of a reaction medium contain relatively few oxidantbubbles. This low volume of oxidant bubbles reduces the amount ofmolecular oxygen available for dissolution into the liquid phase of thereaction medium. Thus, the liquid phase in an unaerated zone of thereaction medium has a relatively low concentration of molecular oxygen.These oxygen-starved, unaerated zones of the reaction medium have atendency to promote undesirable side reactions, rather than the desiredoxidation reaction. For example, when para-xylene is partially oxidizedto form terephthalic acid, insufficient oxygen availability in theliquid phase of the reaction medium can cause the formation ofundesirably high quantities of benzoic acid and coupled aromatic rings,notably including highly undesirable colored molecules known asfluorenones and anthraquinones.

In accordance with one embodiment of the present invention, liquid-phaseoxidation is carried out in a bubble column reactor configured andoperated in a manner such that the volume fraction of the reactionmedium with low gas hold-up values is minimized. This minimization ofunaerated zones can be quantified by theoretically partitioning theentire volume of the reaction medium into 2,000 discrete horizontalslices of uniform volume. With the exception of the highest and lowesthorizontal slices, each horizontal slice is a discrete volume bounded onits sides by the sidewall of the reactor and bounded on its top andbottom by imaginary horizontal planes. The highest horizontal slice isbounded on its bottom by an imaginary horizontal plane and on its top bythe upper surface of the reaction medium. The lowest horizontal slice isbounded on its top by an imaginary horizontal plane and on its bottom bythe lower end of the vessel. Once the reaction medium has beentheoretically partitioned into 2,000 discrete horizontal slices of equalvolume, the time-averaged and volume-averaged gas hold-up of eachhorizontal slice can be determined. When this method of quantifying theamount of unaerated zones is employed, it is preferred for the number ofhorizontal slices having a time-averaged and volume-averaged gas hold-upless than 0.1 to be less than 30, more preferably less than 15, stillmore preferably less than 6; even more preferably less than 4, and mostpreferably less than 2. It is preferred for the number of horizontalslices having a gas hold-up less than 0.2 to be less than 80, morepreferably less than 40, still more preferably less than 20, even morepreferably less than 12, and most preferably less than 5. It ispreferred for the number of horizontal slices having a gas hold-up lessthan 0.3 to be less than 120, more preferably less than 80, still morepreferably less than 40, even more preferably less than 20, and mostpreferably less than 15.

Referring again to FIGS. 1 and 2, it has been discovered thatpositioning oxidant sparger 34 lower in reaction zone 28 providesseveral advantages, including reduction of the amount of unaerated zonesin reaction medium 36. Given a height “H” of reaction medium 36, alength “L” of reaction zone 28, and a maximum diameter “D” of reactionzone 28, it is preferred for a majority (i.e., >50 percent by weight) ofthe oxidant stream to be introduced into reaction zone 28 within about0.025H, 0.022L, and/or 0.25D of lower end 52 of reaction zone 28. Morepreferably, a majority of the oxidant stream is introduced into reactionzone 28 within about 0.02H, 0.018L, and/or 0.2D of lower end 52 ofreaction zone 28. Most preferably, a majority of the oxidant stream isintroduced into reaction zone 28 within 0.015H, 0.013L, and/or 0.15D oflower end 52 of reaction zone 28.

In the embodiment illustrated in FIG. 2, the vertical distance “Y₁”between lower end 52 of reaction zone 28 and the outlet of upper oxidantopenings 68 of oxidant sparger 34 is less than about 0.25H, 0.022L,and/or 0.25D, so that substantially all of the oxidant stream entersreaction zone 28 within about 0.25H, 0.022L, and/or 0.25D of lower end52 of reaction zone 28. More preferably, Y, is less than about 0.02H,0.018L, and/or 0.2D. Most preferably, Y₁ is less than 0.015H, 0.013L,and/or 0.15D, but more than 0.005H, 0.004L, and/or 0.06D. FIG. 2illustrates a tangent line 72 at the location where the bottom edge ofcylindrical main body 46 of vessel shell 22 joins with the top edge ofelliptical lower head 48 of vessel shell 22. Alternatively, lower head48 can be of any shape, including conical, and the tangent line is stilldefined as the bottom edge of cylindrical main body 46. The verticaldistance “Y₂” between tangent line 72 and the top of oxidant sparger 34is preferably at least about 0.0012H, 0.001L, and/or 0.01D; morepreferably at least about 0.005H, 0.004L, and/or 0.05D; and mostpreferably at least 0.01H, 0.008L, and/or 0.1D. The vertical distance“Y₃” between lower end 52 of reaction zone 28 and the outlet of loweroxidant openings 70 of oxidant sparger 34 is preferably less than about0.015H, 0.013L, and/or 0.15D; more preferably less than about 0.012H,0.01L, and/or 0.1D; and most preferably less than 0.01H, 0.008L, and/or0.075D, but more than 0.003H, 0.002L, and/or 0.025D.

In a preferred embodiment of the present invention, the openings thatdischarge the oxidant stream and the feed stream into the reaction zoneare configured so that the amount (by weight) of the oxidant or feedstream discharged from an opening is directly proportional to the openarea of the opening. Thus, for example, if 50 percent of the cumulativeopen area defined by all oxidants openings is located within 0.15D ofthe bottom of the reaction zone, then 50 weight percent of the oxidantstream enters the reaction zone within 0.15D of the bottom of thereaction zone and vice-versa.

In addition to the advantages provided by minimizing unaerated zones(i.e., zones with low gas hold-up) in reaction medium 36, it has beendiscovered that oxidation can be enhanced by maximizing the gas hold-upof the entire reaction medium 36. Reaction medium 36 preferably hastime-averaged and volume-averaged gas hold-up of at least about 0.4,more preferably in the range of from about 0.6 to about 0.9, and mostpreferably in the range of from 0.65 to 0.85. Several physical andoperational attributes of bubble column reactor 20 contribute to thehigh gas hold-up discussed above. For example, for a given reactor sizeand flow of oxidant stream, the high L:D ratio of reaction zone 28yields a lower diameter which increases the superficial velocity inreaction medium 36 which in turn increases gas hold-up. Additionally,the actual diameter of a bubble column and the L:D ratio are known toinfluence the average gas hold-up even for a given constant superficialvelocity. In addition, the minimization of unaerated zones, particularlyin the bottom of reaction zone 28, contributes to an increased gashold-up value. Further, the overhead pressure and mechanicalconfiguration of the bubble column reactor can affect operatingstability at the high superficial velocities and gas hold-up valuesdisclosed herein.

Furthermore, the inventors have discovered the importance of operatingwith an optimized overhead pressure to obtain increased gas hold-up andincreased mass transfer. It might seem that operating with a loweroverhead pressure, which reduces the solubility of molecular oxygenaccording to a Henry's Law effect, would reduce the mass transfer rateof molecular oxygen from gas to liquid. In a mechanically agitatedvessel, such is typically the case because aeration levels and masstransfer rates are dominated by agitator design and overhead pressure.However, in a bubble column reactor according to a preferred embodimentof the present invention, it has been discovered how to use a loweroverhead pressure to cause a given mass of gas-phase oxidant stream tooccupy more volume, increasing the superficial velocity in reactionmedium 36 and in turn increasing the gas hold-up and transfer rate ofmolecular oxygen.

The balance between bubble coalescence and breakup is an extremelycomplicated phenomenon, leading on the one hand to a tendency to foam,which reduces internal circulation rates of the liquid phase and whichmay require very, very large disengaging zones, and on the other hand toa tendency to fewer, very large bubbles that give a lower gas hold-upand lower mass transfer rate from the oxidant stream to the liquidphase. Concerning the liquid phase, its composition, density, viscosityand surface tension, among other factors, are known to interact in avery complicated manner to produce very complicated results even in theabsence of a solid-phase. For example, laboratory investigators havefound it useful to qualify whether “water” is tap water, distilledwater, or de-ionized water, when reporting and evaluating observationsfor even simple water-air bubble columns. For complex mixtures in theliquid phase and for the addition of a solid phase, the degree ofcomplexity rises further. The surface irregularities of individualparticles of solids, the average size of solids, the particle sizedistribution, the amount of solids relative to the liquid phase, and theability of the liquid to wet the surface of the solid, among otherthings, are all important in their interaction with the liquid phase andthe oxidant stream in establishing what bubbling behavior and naturalconvection flow patterns will result.

Thus, the ability of the bubble column reactor to function usefully withthe high superficial velocities and high gas hold-up disclosed hereindepends, for example, on an appropriate selection of: (1) thecomposition of the liquid phase of the reaction medium; (2) the amountand type of precipitated solids, both of which can be adjusted byreaction conditions; (3) the amount of oxidant stream fed to thereactor; (4) the overhead pressure, which affects the volumetric flow ofoxidant stream, the stability of bubbles, and, via the energy balance,the reaction temperature; (5) the reaction temperature itself, whichaffects the fluid properties, the properties of precipitated solids, andthe specific volume of the oxidant stream; and (6) the geometry andmechanical details of the reaction vessel, including the L:D ratio.

Referring again to FIG. 1, it has been discovered that improveddistribution of the oxidizable compound (e.g., para-xylene) in reactionmedium 36 can be provided by introducing the liquid-phase feed streaminto reaction zone 28 at multiple vertically-spaced locations.Preferably, the liquid-phase feed stream is introduced into reactionzone 28 via at least 3 feed openings, more preferably at least 4 feedopenings. As used herein, the term “feed openings” shall denote openingswhere the liquid-phase feed stream is discharged into reaction zone 28for mixing with reaction medium 36. It is preferred for at least 2 ofthe feed openings to be vertically-spaced from one another by at leastabout 0.5D, more preferably at least about 1.5D, and most preferably atleast 3D. However, it is preferred for the highest feed opening to bevertically-spaced from the lowest oxidant opening by not more than about0.75H, 0.65L, and/or 8D; more preferably not more than about 0.5H, 0.4L,and/or 5D; and most preferably not more than 0.4H, 0.35L, and/or 4D.

Although it is desirable to introduce the liquid-phase feed stream atmultiple vertical locations, it has also been discovered that improveddistribution of the oxidizable compound in reaction medium 36 isprovided if the majority of the liquid-phase feed stream is introducedinto the bottom half of reaction medium 36 and/or reaction zone 28.Preferably, at least about 75 weight percent of the liquid-phase feedstream is introduced into the bottom half of reaction medium 36 and/orreaction zone 28. Most preferably, at least 90 weight percent of theliquid-phase feed stream is introduced into the bottom half of reactionmedium 36 and/or reaction zone 28. In addition, it is preferred for atleast about 30 weight percent of the liquid-phase feed stream to beintroduced into reaction zone 28 within about 1.5D of the lowestvertical location where the oxidant stream is introduced into reactionzone 28. This lowest vertical location where the oxidant stream isintroduced into reaction zone 28 is typically at the bottom of oxidantsparger; however, a variety of alternative configurations forintroducing the oxidant stream into reaction zone 28 are contemplated bya preferred embodiment of the present invention. Preferably, at leastabout 50 weight percent of the liquid-phase feed is introduced withinabout 2.5D of the lowest vertical location where the oxidant stream isintroduced into reaction zone 28. Preferably, at least about 75 weightpercent of the liquid-phase feed stream is introduced within about 5D ofthe lowest vertical location where the oxidant stream is introduced intoreaction zone 28.

Each feed opening defines an open area through which the feed isdischarged. It is preferred that at least about 30 percent of thecumulative open area of all the feed inlets is located within about 1.5Dof the lowest vertical location where the oxidant stream is introducedinto reaction zone 28. Preferably, at least about 50 percent of thecumulative open area of all the feed inlets is located within about 2.5Dof the lowest vertical location where the oxidant stream is introducedinto reaction zone 28. Preferably, at least about 75 percent of thecumulative open area of all the feed inlets is located within about 5Dof the lowest vertical location where the oxidant stream is introducedinto reaction zone 28.

Referring again to FIG. 1, in one embodiment of the present invention,feed inlets 32 a,b,c,d are simply a series of vertically-alignedopenings along one side of vessel shell 22. These feed openingspreferably have substantially similar diameters of less than about 7centimeters, more preferably in the range of from about 0.25 to about 5centimeters, and most preferably in the range of from 0.4 to 2centimeters. Bubble column reactor 20 is preferably equipped with asystem for controlling the flow rate of the liquid-phase feed stream outof each feed opening. Such flow control system preferably includes anindividual flow control valve 74 a,b,c,d for each respective feed inlet32 a,b,c,d. In addition, it is preferred for bubble column reactor 20 tobe equipped with a flow control system that allows at least a portion ofthe liquid-phase feed stream to be introduced into reaction zone 28 atan elevated inlet superficial velocity of at least about 2 meters persecond, more preferably at least about 5 meters per second, still morepreferably at least about 6 meters per second, and most preferably inthe range of from 8 to 20 meters per second. As used herein, the term“inlet superficial velocity” denotes the time-averaged volumetric flowrate of the feed stream out of the feed opening divided by the area ofthe feed opening. Preferably, at least about 50 weight percent of thefeed stream is introduced into reaction zone 28 at an elevated inletsuperficial velocity. Most preferably, substantially all the feed streamis introduced into reaction zone 28 at an elevated inlet superficialvelocity.

Referring now to FIGS. 6-7, an alternative system for introducing theliquid-phase feed stream into reaction zone 28 is illustrated. In thisembodiment, the feed stream is introduced into reaction zone 28 at fourdifferent elevations. Each elevation is equipped with a respective feeddistribution system 76 a,b,c,d. Each feed distribution system 76includes a main feed conduit 78 and a manifold 80. Each manifold 80 isprovided with at least two outlets 82,84 coupled to respective insertconduits 86,88, which extend into reaction zone 28 of vessel shell 22.Each insert conduit 86,88 presents a respective feed opening 87,89 fordischarging the feed stream into reaction zone 28. Feed openings 87,89preferably have substantially similar diameters of less than about 7centimeters, more preferably in the range of from about 0.25 to about 5centimeters, and most preferably in the range of from 0.4 to 2centimeters. It is preferred for feed openings 87,89 of each feeddistribution system 76 a,b,c,d to be diametrically opposed so as tointroduce the feed stream into reaction zone 28 in opposite directions.Further, it is preferred for the diametrically opposed feed openings86,88 of adjacent feed distribution systems 76 to be oriented at 90degrees of rotation relative to one another. In operation, theliquid-phase feed stream is charged to main feed conduit 78 andsubsequently enters manifold 80. Manifold 80 distributes the feed streamevenly for simultaneous introduction on opposite sides of reactor 20 viafeed openings 87,89.

FIG. 8 illustrates an alternative configuration wherein each feeddistribution system 76 is equipped with bayonet tubes 90,92 rather thaninsert conduits 86,88 (shown in FIG. 7). Bayonet tubes 90,92 projectinto reaction zone 28 and include a plurality of small feed openings94,96 for discharging the liquid-phase feed into reaction zone 28. It ispreferred for the small feed openings 94,96 of bayonet tubes 90,92 tohave substantially the same diameters of less than about 50 millimeters,more preferably about 2 to about 25 millimeters, and most preferably 4to 15 millimeters.

FIGS. 9-11 illustrate an alternative feed distribution system 100. Feeddistribution system 100 introduces the liquid-phase feed stream at aplurality of vertically-spaced and laterally-spaced locations withoutrequiring multiple penetrations of the sidewall of bubble column reactor20. Feed introduction system 100 generally includes a single inletconduit 102, a header 104, a plurality of upright distribution tubes106, a lateral support mechanism 108, and a vertical support mechanism110. Inlet conduit 102 penetrates the sidewall of main body 46 of vesselshell 22. Inlet conduit 102 is fluidly coupled to header 104. Header 104distributes the feed stream received from inlet conduit 102 evenly amongupright distribution tubes 106. Each distribution tube 106 has aplurality of vertically-spaced feed openings 112 a,b,c,d for dischargingthe feed stream into reaction zone 28. Lateral support mechanism 108 iscoupled to each distribution tube 106 and inhibits relative lateralmovement of distribution tubes 106. Vertical support mechanism 110 ispreferably coupled to lateral support mechanism 108 and to the top ofoxidant sparger 34. Vertical support mechanism 110 substantiallyinhibits vertical movement of distribution tubes 106 in reaction zone28. It is preferred for feed openings 112 to have substantially the samediameters of less than about 50 millimeters, more preferably about 2 toabout 25 millimeters, and most preferably 4 to 15 millimeters. Thevertical spacing of feed openings 112 of feed distribution system 100illustrated in FIGS. 9-11 can be substantially the same as describedabove with reference to the feed distribution system of FIG. 1.

It has been discovered that the flow patterns of the reaction medium inmany bubble column reactors can permit uneven azimuthal distribution ofthe oxidizable compound in the reaction medium, especially when theoxidizable compound is primarily introduced along one side of thereaction medium. As used herein, the term “azimuthal” shall denote anangle or spacing around the upright axis of elongation of the reactionzone. As used herein, “upright” shall mean within 45° of vertical. Inone embodiment of the present invention, the feed stream containing theoxidizable compound (e.g., para-xylene) is introduced into the reactionzone via a plurality of azimuthally-spaced feed openings. Theseazimuthally-spaced feed openings can help prevent regions of excessivelyhigh and excessively low oxidizable compound concentrations in thereaction medium. The various feed introduction systems illustrated inFIGS. 6-11 are examples of systems that provide proper azimuthal spacingof feed openings.

Referring again to FIG. 7, in order to quantify the azimuthally-spacedintroduction of the liquid-phase feed stream into the reaction medium,the reaction medium can be theoretically partitioned into four uprightazimuthal quadrants “Q₁,Q₂,Q₃,Q₄” of approximately equal volume. Theseazimuthal quadrants “Q₁,Q₂,Q₃,Q₄” are defined by a pair of imaginaryintersecting perpendicular vertical planes “P₁,P₂” extending beyond themaximum vertical dimension and maximum radial dimension of the reactionmedium. When the reaction medium is contained in a cylindrical vessel,the line of intersection of the imaginary intersecting vertical planesP₁,P₂ will be approximately coincident with the vertical centerline ofthe cylinder, and each azimuthal quadrant Q₁,Q₂,Q₃,Q₄ will be agenerally wedge-shaped vertical volume having a height equal to theheight of the reaction medium. It is preferred for a substantial portionof the oxidizable compound to be discharged into the reaction medium viafeed openings located in at least two different azimuthal quadrants.

In a preferred embodiment of the present invention, not more than about80 weight percent of the oxidizable compound is discharged into thereaction medium through feed openings that can be located in a singleazimuthal quadrant. More preferably, not more than about 60 weightpercent of the oxidizable compound is discharged into the reactionmedium through feed openings that can be located in a single azimuthalquadrant. Most preferably, not more than 40 weight percent of theoxidizable compound is discharged into the reaction medium through feedopenings that can be located in a single azimuthal quadrant. Theseparameters for azimuthal distribution of the oxidizable compound aremeasured when the azimuthal quadrants are azimuthally oriented such thatthe maximum possible amount of oxidizable compound is being dischargedinto one of the azimuthal quadrants. For example, if the entire feedstream is discharged into the reaction medium via two feed openings thatare azimuthally spaced from one another by 89 degrees, for purposes ofdetermining azimuthal distribution in four azimuthal quadrants, 100weight percent of the feed stream is discharged into the reaction mediumin a single azimuthal quadrant because the azimuthal quadrants can beazimuthally oriented in such a manner that both of the feed openings arelocated in a single azimuthal quadrant.

In addition to the advantages associated with the properazimuthal-spacing of the feed openings, it has also been discovered thatproper radial spacing of the feed openings in a bubble column reactorcan also be important. It is preferred for a substantial portion of theoxidizable compound introduced into the reaction medium to be dischargedvia feed openings that are radially spaced inwardly from the sidewall ofthe vessel. Thus, in one embodiment of the present invention, asubstantial portion of the oxidizable compound enters the reaction zonevia feed openings located in a “preferred radial feed zone” that isspaced inwardly from the upright sidewalls defining the reaction zone.

Referring again to FIG. 7, the preferred radial feed zone “FZ” can takethe shape of a theoretical upright cylinder centered in reaction zone 28and having an outer diameter “D_(O)” of 0.9D, where “D” is the diameterof reaction zone 28. Thus, an outer annulus “OA” having a thickness of0.05D is defined between the preferred radial feed zone FZ and theinside of the sidewall defining reaction zone 28. It is preferred forlittle or none of the oxidizable compound to be introduced into reactionzone 28 via feed openings located in this outer annulus OA.

In another embodiment, it is preferred for little or none of theoxidizable compound to be introduced into the center of reaction zone28. Thus, as illustrated in FIG. 8, the preferred radial feed zone FZcan take the shape of a theoretical upright annulus centered in reactionzone 28, having an outer diameter D_(O) of 0.9D, and having an innerdiameter D_(I) of 0.2D. Thus, in this embodiment, an inner cylinder IChaving a diameter of 0.2D is “cut out” of the center of the preferredradial feed zone FZ. It is preferred for little or none of theoxidizable compound to be introduced into reaction zone 28 via feedopenings located in this inner cylinder IC.

In a preferred embodiment of the present invention, a substantialportion of the oxidizable compound is introduced into reaction medium 36via feed openings located in the preferred radial feed zone, regardlessof whether the preferred radial feed zone has the cylindrical or annularshape described above. More preferably, at least about 25 weight percentof the oxidizable compound is discharged into reaction medium 36 viafeed openings located in the preferred radial feed zone. Still morepreferably, at least about 50 weight percent of the oxidizable compoundis discharged into reaction medium 36 via feed openings located in thepreferred radial feed zone. Most preferably, at least 75 weight percentof the oxidizable compound is discharged into reaction medium 36 viafeed openings located in the preferred radial feed zone.

Although the theoretical azimuthal quadrants and theoretical preferredradial feed zone illustrated in FIGS. 7 and 8 are described withreference to the distribution of the liquid-phase feed stream, it hasbeen discovered that proper azimuthal and radial distribution of thegas-phase oxidant stream can also provide certain advantages. Thus, inone embodiment of the present invention, the description of theazimuthal and radial distribution of the liquid-phase feed stream,provided above, also applies to the manner in which the gas-phaseoxidant stream is introduced into the reaction medium 36.

Referring now to FIGS. 12-15, an alternative oxidant sparger 200 isillustrated as generally comprising a ring member 202 and a pair ofoxidant entry conduits 204,206. Oxidant sparger 200 of FIGS. 12-15 issimilar to oxidant sparger 34 of FIGS. 1-11 with the following threeprimary differences: (1) oxidant sparger 200 does not include a diagonalcross-member; (2) the upper portion of ring member 202 has no openingsfor discharging the oxidant in an upward direction; and (3) oxidantsparger 200 has many more openings in the lower portion of ring member202.

As perhaps best illustrated in FIGS. 14 and 15, the bottom portion ofoxidant sparger ring 202 presents a plurality of oxidant openings 208.Oxidant openings 208 are preferably configured such that at least about1 percent of the total open area defined by oxidant openings 208 islocated below the centerline 210 (FIG. 15) of ring member 202, wherecenterline 210 is located at the elevation of the volumetric centroid ofring member 202. More preferably, at least about 5 percent of the totalopen area defined by all oxidant openings 208 is located belowcenterline 210, with at least about 2 percent of the total open areabeing defined by openings 208 that discharge the oxidant stream in agenerally downward direction within about 30 degrees of vertical. Stillmore preferably, at least about 20 percent of the total open areadefined by all oxidant openings 208 is located below centerline 210,with at least about 10 percent of the total open area being defined byopenings 208 that discharge the oxidant stream in a generally downwarddirection within 30 degrees of vertical. Most preferably, at least about75 percent of the total open area defined by all oxidant openings 208 islocated below centerline 210, with at least about 40 percent of thetotal open area being defined by openings 208 that discharge the oxidantstream in a generally downward direction within 30 degrees of vertical.The fraction of the total open area defined by all oxidant openings 208that are located above centerline 210 is preferably less than about 75percent, more preferably less than about 50 percent, still morepreferably less than about 25 percent, and most preferably less than 5percent.

As illustrated in FIGS. 14 and 15, oxidant openings 208 include downwardopenings 208 a and skewed openings 208 b. Downward openings 208 a areconfigured to discharge the oxidant stream generally downwardly at anangle within about 30 degrees of vertical, more preferably within about15 degrees of vertical, and most preferably within 5 degrees ofvertical. Skewed openings 208 b are configured to discharge the oxidantstream generally outwardly and downwardly at an angle “A” that is in therange of from about 15 to about 75 degrees from vertical, morepreferably angle A is in the range of from about 30 to about 60 degreesfrom vertical, and most preferably angle A is in the range of from 40 to50 degrees from vertical.

It is preferred for substantially all oxidant openings 208 to haveapproximately the same diameter. The diameter of oxidant openings 208 ispreferably in the range of from about 2 to about 300 millimeters, morepreferably in the range of from about 4 to about 120 millimeters, andmost preferably in the range of from 8 to 60 millimeters. The totalnumber of oxidant openings 208 in ring member 202 is selected to meetthe low pressure drop criteria detailed below. Preferably, the totalnumber of oxidant openings 208 formed in ring member 202 is at leastabout 10, more preferably the total number of oxidant openings 208 is inthe range of from about 20 to about 200, and most preferably the totalnumber of oxidant openings 208 is in the range of from 40 to 100.

Although FIGS. 12-15 illustrate a very specific configuration foroxidant sparger 200, it is now noted that a variety of oxidant spargerconfigurations can be employed to achieve the advantages describedherein. For example, the oxidant sparger does not necessarily need tohave the octagonal ring member configuration illustrated in FIGS. 12-13.Rather, it is possible for the oxidant sparger to be formed of anyconfiguration of flow conduit(s) that employs a plurality ofspaced-apart openings for discharging the oxidant stream. The size,number, and discharge direction of the oxidant openings in the flowconduit are preferably within the ranges stated above. Further, theoxidant sparger is preferably configured to provide the azimuthal andradial distribution of molecular oxygen described above.

Regardless of the specific configuration of the oxidant sparger, it ispreferred for the oxidant sparger to be physically configured andoperated in a manner that minimizes the pressure drop associated withdischarging the oxidant stream out of the flow conduit(s), through theoxidant openings, and into the reaction zone. Such pressure drop iscalculated as the time-averaged static pressure of the oxidant streaminside the flow conduit at oxidant inlets 66 a,b of the oxidant spargerminus the time-averaged static pressure in the reaction zone at theelevation where one-half of the oxidant stream is introduced above thatvertical location and one-half of the oxidant stream is introduced belowthat vertical location. In a preferred embodiment of the presentinvention, the time-averaged pressure drop associated with dischargingthe oxidant stream from the oxidant sparger is less than about 0.3megapascal (MPa), more preferably less than about 0.2 MPa, still morepreferably less than about 0.1 MPa, and most preferably less than 0.05MPa. Under the preferred operating conditions of the bubble columnreactor described herein, the pressure of the oxidant stream inside theflow conduit(s) of the oxidant sparger is preferably in the range offrom about 0.35 to about 1 MPa, more preferably in the range of fromabout 0.45 to about 0.85 MPa, and most preferably in the range of from0.5 to 0.7 MPa.

As alluded to earlier with reference to the oxidant spargerconfiguration illustrated in FIGS. 2-5, it may be desirable tocontinuously or periodically flush the oxidant sparger with a liquid(e.g., acetic acid, water, and/or para-xylene) to prevent fouling of theoxidant sparger with solids. When such a liquid flush is employed, it ispreferred for an effective amount of the liquid (i.e., not just theminor amount of liquid droplets that might naturally be present in theoxidant stream) to be passed through the oxidant sparger and out of theoxidant openings for at least one period of more than one minute eachday. When a liquid is continuously or periodically discharged from theoxidant sparger, it is preferred for the time-averaged ratio of the massflow rate of the liquid through the oxidant sparger to the mass flowrate of the molecular oxygen through the oxidant sparger to be in therange of from about 0.05:1 to about 30:1, or in the range of from about0.1:1 to about 2:1, or even in the range of from 0.2:1 to 1:1.

In one embodiment of the present invention, a significant portion of theoxidizable compound (e.g., para-xylene) can be introduced into thereaction zone through the oxidant sparger. In such a configuration, itis preferred for the oxidizable compound and the molecular oxygen to bedischarged from the oxidant sparger through the same openings in theoxidant sparger. As noted above, the oxidizable compound is typically aliquid at STP. Therefore, in this embodiment, a two-phase stream may bedischarged from the oxidant sparger, with the liquid phase comprisingthe oxidizable compound and the gas phase comprising the molecularoxygen. It should be recognized, however, that at least a portion of theoxidizable compound may be in a gaseous state when discharged from theoxidant sparger. In one embodiment, the liquid phase discharged from theoxidant sparger is formed predominately of the oxidizable compound. Inanother embodiment, the liquid phase discharged from the oxidant spargerhas substantially the same composition as the feed stream, describedabove. When the liquid phase discharged from the oxidant sparger hassubstantially the same composition as the feed stream, such liquid phasemay comprise a solvent and/or a catalyst system in the amounts andratios described above with reference to the composition of the feedstream.

In one embodiment of the present invention, it is preferred for at leastabout 10 weight percent of all the oxidizable compound introduced intothe reaction zone to be introduced via the oxidant sparger, morepreferably at least about 40 weight percent of the oxidizable compoundis introduced into the reaction zone via the oxidant sparger, and mostpreferably at least 80 weight percent of the oxidizable compound isintroduced into the reaction zone via the oxidant sparger. When all orpart of the oxidizable compound is introduced into the reaction zone viathe oxidant sparger, it is preferred for at least about 10 weightpercent of all the molecular oxygen introduced into the reaction zone tobe introduced via the same oxidant sparger, more preferably at leastabout 40 weight percent of the oxidizable compound is introduced intothe reaction zone via the same oxidant sparger, and most preferably atleast 80 weight percent of the oxidizable compound is introduced intothe reaction zone via the same oxidant sparger. When a significantportion of the oxidizable compound is introduced into the reaction zonevia the oxidant sparger, it is preferred for one or more temperaturesensing devices (e.g., thermocouples) to be disposed in the oxidantsparger. These temperature sensors can be employed to help to make surethe temperature in the oxidant sparger does not become dangerously high.

Referring now to FIGS. 16-18, bubble column reactor 20 is illustrated asincluding an internal deaeration vessel 300 disposed in the bottom ofreaction zone 28 near slurry outlet 38. It has been discovered thatimpurity-forming side reactions occur at a relatively high rate duringdeaeration of reaction medium 36. As used herein, “deaeration” shalldenote the disengagement of a gas phase from multi-phase reactionmedium. When reaction medium 36 is highly aerated (>0.3 gas hold-up),impurity formation is minimal. When reaction medium 36 is highlyunaerated (<0.01 gas hold-up), impurity formation is also minimal.However, when reaction medium is partially-aerated (0.01-0.3 gashold-up), undesirable side reactions are promoted and increasedimpurities are generated. Deaeration vessel 300 addresses this and otherproblems by minimizing the volume of reaction medium 36 in apartially-aerated stated, and by minimizing the time it takes todeaerate reaction medium 36. A substantially deaerated slurry isproduced from the bottom of deaeration vessel 300 and exits reactor 20via slurry outlet 38. The substantially deaerated slurry preferablycontains less than about 5 volume percent gas phase, more preferablyless than about 2 volume percent gas phase, and most preferably lessthan 1 volume percent gas phase.

In FIG. 16, bubble column reactor 20 is illustrated as including a levelcontroller 302 and a flow control valve 304. Level controller 302 andflow control valve 304 cooperate to maintain reaction medium 36 at asubstantially constant elevation in reaction zone 28. Level controller302 is operable to sense (e.g., by differential pressure level sensingor by nuclear level sensing) the elevation of upper surface 44 ofreaction medium 36 and generate a control signal 306 responsive to theelevation of reaction medium 36. Flow control valve 304 receives controlsignal 306 and adjusts the flow rate of a slurry through a slurry outletconduit 308. Thus, the flow rate of the slurry out of slurry outlet 38can vary between a maximum slurry volumetric flow rate (F_(max)) whenthe elevation of reaction medium 36 is too high and a minimum slurryvolumetric flow rate (F_(min)) when the elevation of reaction medium 36is too low.

In order to remove solid-phase oxidation product from reaction zone 28,a portion must first pass through deaeration vessel 300. Deaerationvessel 300 provides a low-turbulence internal volume that permits thegas phase of reaction medium 36 to naturally rise out of the liquid andsolid phases of reaction medium 36 as the liquid and solids flowdownwardly toward slurry outlet 38. The rising of the gas phase out ofthe liquid and solid phases is caused by the natural upward buoyancy ofthe gas phase in the liquid and solid phases. When deaeration vessel 300is employed, the transitioning of reaction medium 36 from afully-aerated, three-phase medium to a fully-deaerated, two-phase slurryis quick and efficient.

Referring now to FIGS. 17 and 18, deaeration vessel 300 includes agenerally upright sidewall 308 defining a deaeration zone 312therebetween. Preferably, sidewall 308 extends upwardly within about 30degrees of vertical, more preferably within about 10 degrees ofvertical. Most preferably, sidewall 308 is substantially vertical.Deaeration zone 312 is separate from reaction zone 28 and has height “h”and a diameter “d.” An upper end 310 of sidewall 308 is open so as toreceive reaction medium from reaction zone 28 into internal volume 312.The lower end of sidewall 308 is fluidly coupled to slurry outlet 38 viaa transition section 314. In certain instances, such as when the openingof slurry outlet 38 is large or when the diameter “d” of sidewall 308 issmall, transition section 314 can be eliminated. As perhaps bestillustrated in FIG. 18, deaeration vessel 300 can also include a vortexbreaker 316 disposed in deaeration zone 312. Vortex breaker 316 can beany structure operable to inhibit the formation of vortices as the solidand liquid phases flow downwardly towards slurry outlet 38.

In order to permit proper disengagement of the gas phase from the solidand liquid phases in deaeration vessel 300, the height “h” andhorizontal cross-sectional area of internal deaeration zone 312 arecarefully selected. The height “h” and horizontal cross-sectional areaof internal deaeration zone 312 should provide sufficient distance andtime so that even when the maximum amount of slurry is being withdrawn(i.e., when slurry is being withdrawn at F_(max)), substantially all ofthe gas bubble volume can rise out of the solid and liquid phases beforethe gas bubbles reach the bottom outlet of deaeration vessel 300. Thus,it is preferred for the cross-sectional area of deaeration zone 312 tobe such that the maximum downward velocity (V_(dmax)) of the liquid andsolid phases through deaeration zone 312 is substantially less than thenatural rise velocity (V_(u)) of the gas phase bubbles through theliquid and solid phases. The maximum downward velocity (V_(dmax)) of theliquid and solid phases through deaeration zone 312 occurs at themaximum slurry volumetric flow rate (F_(max)), discussed above. Thenatural rise velocity (V_(u)) of the gas bubbles through the liquid andsolid phases varies depending on the size of the bubbles; however, thenatural rise velocity (V_(u0.5)) of 0.5 centimeter diameter gas bubblesthrough the liquid and solid phases can be used as a cut-off valuebecause substantially all of the bubble volume initially in reactionmedium 36 will be greater than 0.5 centimeters. Preferably, thecross-sectional area of deaeration zone 312 is such that V_(dmax) isless than about 75 percent of V_(u0.5), more preferably V_(dmax) is lessthan about 40 percent of V_(u0.5), most preferably V_(dmax) is less than20 percent of V_(u0.5).

The downward velocity of the liquid and solid phases in deaeration zone312 of deaeration vessel 300 is calculated as the volumetric flow rateof the deaerated slurry through slurry outlet 38 divided by the minimumcross-sectional area of deaeration zone 312. The downward velocity ofthe liquid and solid phases in deaeration zone 312 of deaeration vessel300 is preferably less than about 50 centimeters per second, morepreferably less than about 30 centimeters per second, and mostpreferably less than 10 centimeters per second.

It is now noted that although upright sidewall 308 of deaeration vessel300 is illustrated as having a cylindrical configuration, sidewall 308could comprise a plurality of sidewalls that form a variety ofconfigurations (e.g., triangular, square, or oval), so long as the wallsdefines an internal volume having an appropriate volume, cross-sectionalarea, width “d”, and height “h”. In a preferred embodiment of thepresent invention, “d” is in the range of from about 0.2 to about 2meters, more preferably in the range of from about 0.3 to about 1.5meters, and most preferably in the range of from 0.4 to 1.2 meters. In apreferred embodiment of the present invention, “h” is in the range offrom about 0.3 meters to about 5 meters, more preferably in the range offrom about 0.5 to about 3 meters, and most preferably in the range offrom 0.75 to 2 meters.

In a preferred embodiment of the present invention, sidewall 308 issubstantially vertical so that the horizontal cross-sectional area ofdeaeration zone 312 is substantially constant along the entire height“h” of deaeration zone 312. Preferably, the maximum horizontalcross-sectional area of deaeration zone 312 is less than about 25percent of the maximum horizontal cross-sectional area of reaction zone28. More preferably, the maximum horizontal cross-sectional area ofdeaeration zone 312 is in the range of from about 0.1 to about 10percent of the maximum horizontal cross-sectional area of reaction zone28. Most preferably, the maximum horizontal cross-sectional area ofdeaeration zone 312 is in the range of from 0.25 to 4 percent of themaximum horizontal cross-sectional area of reaction zone 28. Preferably,the maximum horizontal cross-sectional area of deaeration zone 312 is inthe range of from about 0.02 to about 3 square meters, more preferablyin the range of from about 0.05 to about 2 square meters, and mostpreferably in the range of from 0.1 to 1.2 square meters. The volume ofdeaeration zone 312 is preferably less than about 5 percent of the totalvolume of reaction medium 36 or reaction zone 28. More preferably, thevolume of deaeration zone 312 is in the range of from about 0.01 toabout 2 percent of the total volume of reaction medium 36 or reactionzone 28. Most preferably, the volume of deaeration zone 312 is in therange of from 0.05 to about 1 percent of the total volume of reactionmedium 36 or reaction zone 28. The volume of deaeration zone 312 ispreferably less than about 2 cubic meters, more preferably in the rangeof from about 0.01 to about 1 cubic meters, and most preferably in therange of from 0.05 to 0.5 cubic meters.

Turning now to FIG. 19, bubble column reactor 20 is illustrated asincluding an external deaeration vessel 400. In this configuration,aerated reaction medium 36 is withdrawn from reaction zone 28 via anelevated opening in the side of vessel shell 22. The withdrawn aeratedmedium is transported to external deaeration vessel 400 via an outletconduit 402 for disengagement of the gas phase from the solid and liquidphases. The disengaged gas phase exits deaeration vessel 400 via conduit404, while the substantially deaerated slurry exits deaeration vessel400 via conduit 406.

In FIG. 19, outlet conduit 402 is shown as being approximately straight,horizontal, and orthogonal to vessel shell 22. This is merely oneconvenient configuration; and outlet conduit 402 may be otherwise in anyrespect, providing that it usefully connects bubble column reactor 20with external deaeration vessel 400. Turning to conduit 404, it isuseful for this conduit to connect at or near the top deaeration vessel400 in order to control safety issues relating to a stagnant gas pocketcontaining oxidizable compound and oxidant. Furthermore, conduits 402and 404 may usefully comprise means of flow isolation, such as valves.

When reaction medium 36 is withdrawn from reactor 20 via an elevatedoutlet, as shown in FIG. 19, it is preferred for bubble column reactor20 to be equipped with a lower outlet 408 near the bottom 52 of reactionzone 28. Lower outlet 408 and a lower conduit 410, coupled thereto, canbe used to de-inventory (i.e., empty) reactor 20 during shutdowns.Preferably, one or more lower outlet 408 is provided in the bottomone-third of the height of reaction medium 36, more preferably in thebottom one-fourth of reaction medium 36, and most preferably at thelowest point of reaction zone 28.

With the elevated slurry withdrawal and deaeration system shown in FIG.19, lower conduit 410 and outlet 408 are not used to withdraw slurryfrom reaction zone 28 during oxidation. It is known in the art thatsolids tend to settle by gravity forces in unaerated and otherwiseunagitated portions of the slurry, including in stagnant flow conduits.Furthermore, the settled solids (e.g., terephthalic acid) can tend tosolidify into large agglomerates by continuing precipitation and/orcrystalline reorganization. Thus, in order to avoid plugging of lowerflow conduit 410, a fraction of the deaerated slurry from the bottom ofdeaeration vessel 400 can be used to continuously or intermittentlyflush lower conduit 410 during normal operation of reactor 20. Apreferred means of providing such a slurry flush to conduit 410 is toperiodically open a valve 412 in conduit 410 and allow a fraction of thedeaerated slurry to flow through conduit 410 and into reaction zone 28via lower opening 408. Even when valve 412 is fully or partially open,only a fraction of the deaerated slurry flows through lower conduit 410and back into reaction zone 28. The remaining fraction of the deaeratedslurry not used to flush lower conduit 410 is carried via conduit 414away from reactor 20 for further downstream processing (e.g.,purification).

During normal operation of bubble column reactor 20 over a substantiallength of time (e.g., >100 hours), it is preferred for the amount ofdeaerated slurry used to flush lower conduit 410 to be less than 50percent by weight of the total deaerated slurry produced from the bottomof deaeration vessel 400, more preferably less than about 20 percent byweight, and most preferably less than 5 percent by weight. Further, itis preferred that over a substantial length of time the average massflow rate of deaerated slurry used to flush lower conduit 410 is lessthan about 4 times the average mass flow rate of the oxidizable compoundinto reaction zone 28, more preferably less than about 2 times theaverage mass flow rate of the oxidizable compound into reaction zone 28,still more preferably less than the average mass flow rate of theoxidizable compound into reaction zone 28, and most preferably less than0.5 times the average mass flow rate of the oxidizable compound intoreaction zone 28.

Referring again to FIG. 19, deaeration vessel 400 includes asubstantially upright, preferably cylindrical sidewall 416 defining adeaeration zone 418. Deaeration zone 418 has a diameter “d” and height“h.” Height “h” is measured as the vertical distance between thelocation where the aerated reaction medium enters deaeration vessel 400and the bottom of sidewall 416. The height “h”, diameter “d”, area, andvolume of deaeration zone 418 is preferably substantially the same asdescribed above with reference to deaeration zone 312 of deaerationvessel 300 illustrated in FIGS. 16-18. In addition, deaeration vessel400 includes an upper section 420 formed by extending sidewall 416 abovedeaeration zone 418. Upper section 420 of deaeration vessel 400 may beof any height, though it preferably extends upwardly to or above thelevel of reaction medium 36 in reaction zone 28. Upper section 420ensures that the gas phase has room to properly disengage from theliquid and solid phases before exiting deaeration vessel 400 via conduit404. It is now noted that although conduit 404 is illustrated asreturning the disengaged gas phase to the disengagement zone of reactor20, conduit 404 could alternatively be coupled to vessel shell 22 at anyelevation above outlet conduit 402. Optionally, conduit 404 could becoupled to gas outlet conduit 40 so that the disengaged gas phase fromdeaeration vessel 400 is combined with the removed overhead vapor streamin conduit 40 and sent downstream for further processing.

Turning now to FIG. 20, bubble column reactor 20 is illustrated asincluding a hybrid internal-external deaeration vessel 500. In thisconfiguration, a portion of reaction medium 36 is withdrawn fromreaction zone 28 through a relatively large elevated opening 502 in thesidewall of vessel shell 22. The withdrawn reaction medium 36 is thentransported through an elbow conduit 504 of relatively large diameterand enters the top of deaeration vessel 500. In FIG. 20, elbow conduit504 is shown as connecting orthogonally to the sidewall of vessel shell22 and as comprising a smooth turn through an angle of about 90 degrees.This is merely one convenient configuration; and elbow conduit 504 maybe otherwise in any respect, providing that it usefully connects bubblecolumn reactor 20 with external deaeration vessel 500, as described.Furthermore, elbow conduit 504 may usefully comprise means of flowisolation, such as valves.

In deaeration vessel 500, the gas phase moves upwardly, while the solidand liquid phases move downwardly. The upwardly moving gas phase canre-enter elbow conduit 504 and then escape through opening 502 back intoreaction zone 28. Thus, a counter-current flow of the entering reactionmedium 36 and the exiting disengaged gas can occur at opening 502. Thedeaerated slurry exits deaeration vessel 500 via conduit 506. Deaerationvessel 500 includes a substantially upright, preferably cylindricalsidewall 508 defining a deaeration zone 510. Deaeration zone 510 has aheight “h” and a diameter “d.” It is preferred for elevated opening 502and elbow conduit 504 to have a diameter the same as, or greater than,the diameter “d” of deaeration zone 510. The height “h”, diameter “d”,area, and volume of deaeration zone 510 are preferably substantially thesame as described above with reference to deaeration zone 312 ofdeaeration vessel 300 illustrated in FIGS. 16-18.

FIGS. 19 and 20 illustrate an embodiment of bubble column reactor 20where the solid product (e.g., crude terephthalic acid) produced inreaction zone 28 is withdrawn from reaction zone 28 via an elevatedoutlet. Withdrawing aerated reaction medium 36 from an elevated locationabove the bottom of bubble column reactor 20 can help avoid accumulationand stagnation of poorly aerated reaction medium 36 at the bottom 52 ofreaction zone 28. According to other aspects of the present invention,the concentrations of oxygen and the oxidizable compound (e.g.,para-xylene) in the reaction medium 36 near the top of reaction medium36 are preferably lower than near the bottom. Thus, withdrawing reactionmedium 36 at an elevated location can increase yield by lowering theamount of unreacted reactants withdrawn from reactor 20. Also, thetemperature of reaction medium 36 varies significantly in the verticaldirection when bubble column reactor 20 is operated with the high STRand the gradients of chemical composition as disclosed herein. Undersuch conditions, the temperature of reaction medium 36 will typicallyhave local minima near the lower end and the upper end of reaction zone28. Near the lower end, the minimum relates to the evaporation ofsolvent near where all or part of the oxidant is admitted. Near theupper end, the minimum is again due to evaporation of solvent, thoughhere due to declining pressure within the reaction medium. In addition,other local minima may occur in between the upper and lower endswherever additional feed or oxidant is admitted to the reaction medium.Thus, there exist one or more temperature maxima, driven by theexothermic heat of oxidation reactions, between the lower end and upperend of reaction zone 28. Withdrawing reaction medium 36 at an elevatedlocation of higher temperature can be particularly advantageous whendownstream processing occurs at higher temperatures, because energycosts associated with heating the withdrawn medium for downstreamprocessing are reduced.

Thus, in a preferred embodiment of the present invention and especiallywhen downstream processing occurs at higher temperatures, reactionmedium 36 is withdrawn from bubble column reactor 20 via an elevatedoutlet(s) positioned above the location(s) where at least 50 weightpercent of the liquid-phase feed stream and/or the gas-phase oxidantstream enter reaction zone 28. More preferably, reaction medium 36 iswithdrawn from bubble column reactor 20 via an elevated outlet(s)positioned above the location(s) where substantially all of theliquid-phase feed stream and/or the gas-phase oxidant stream enterreaction zone 28. Preferably, at least 50 weight percent of thesolid-phase and liquid-phase components withdrawn from bubble columnreactor 20 are withdrawn via an elevated outlet(s). More preferably,substantially all of the solid-phase and liquid-phase componentswithdrawn from bubble column reactor 20 are withdrawn via an elevatedoutlet(s). Preferably, the elevated outlet(s) is located at least about1D above lower end 52 of reaction zone 28. More preferably, the elevatedoutlet(s) is located at least about 2D above lower end 52 of reactionzone 28. Most preferably, the elevated outlet(s) is located at least 3Dabove lower end 52 of reaction zone 28. Given a height “H” of reactionmedium 36, it is preferred for the elevated outlet(s) to be verticallylocated between about 0.2H and about 0.8H, more preferably between about0.3H and about 0.7H, and most preferably between 0.4H and 0.6H.Furthermore, it is preferred that the temperature of reaction medium 36at an elevated outlet from reaction zone 28 is at least 1° C. greaterthan the temperature of reaction medium 36 at lower end 52 of reactionzone 28. More preferably, the temperature of reaction medium 36 at theelevated outlet of reaction zone 28 is in the range of from about 1.5 toabout 16° C. hotter than the temperature of reaction medium 36 at lowerend 52 of reaction zone 28. Most preferably, the temperature of reactionmedium 36 at the elevated outlet of reaction zone 28 is in the range offrom 2 to 12° C. hotter than the temperature of reaction medium 36 atlower end 52 of reaction zone 28.

Referring now to FIG. 21, bubble column reactor 20 is illustrated asincluding an alternative hybrid deaeration vessel 600 positioned at thebottom of reactor 20. In this configuration, aerated reaction medium 36is withdrawn from reaction zone 28 through a relatively large opening602 in the lower end 52 of vessel shell 22. Opening 602 defines the openupper end of deaeration vessel 600. In deaeration vessel 600, the gasphase moves upwardly, while the solid and liquid phases move downwardly.The upwardly moving gas phase can re-enter reaction zone 28 throughopening 602. Thus, a counter-current flow of the entering reactionmedium 36 and the exiting disengaged gas can occur at opening 602. Thedeaerated slurry exits deaeration vessel 600 via conduit 604. Deaerationvessel 600 includes a substantially upright, preferably cylindricalsidewall 606 defining a deaeration zone 608. Deaeration zone 608 has aheight “h” and a diameter “d.” It is preferred for opening 602 to have adiameter the same as, or greater than, the diameter “d” of deaerationzone 608. The height “h”, diameter “d”, area, and volume of deaerationzone 608 are preferably substantially the same as described above withreference to deaeration zone 312 of deaeration vessel 300 illustrated inFIGS. 16-18.

Referring now to FIG. 22, bubble column reactor 20 of FIG. 21 isillustrated as including an alternative oxidant sparger 620. Oxidantsparger 620 includes a ring member 622 and a pair of inlet conduits624,626. Ring member 622 preferably has substantially the sameconfiguration as ring member 202, described above with reference toFIGS. 12-15. Inlet conduits 624,626 extend upwardly through openings inlower head 48 of vessel shell 22 and provide the oxidant stream to ringmember 622.

Referring now to FIG. 23, bubble column reactor 20 of FIG. 21 isillustrated as including a spargerless means for introducing the oxidantstream into reaction zone 28. In the configuration of FIG. 23, theoxidant stream is provided, to reactor 20 via oxidant conduits 630,632.Oxidant conduits 630,632 are coupled to respective oxidant openings634,636 in lower head 48 of vessel shell 22. The oxidant stream isintroduced directly into reaction zone 28 via oxidant openings 634,636.Optional impingement plates 638,640 can be provided to deflect the flowof the oxidant stream once it has initially entered reaction zone 28.

As mentioned above, it is preferred for the oxidation reactor to beconfigured and operated in a manner that avoids zones of highconcentration of oxidizable compound in the reaction medium because suchzones can lead to the formation of impurities. One way to improveinitial dispersion of the oxidizable compound (e.g., para-xylene) in thereaction medium is by diluting the oxidizable compound with a liquid.The liquid used to dilute the oxidizable compound can originate from aportion of the reaction medium located a substantial distance from thelocation(s) where the oxidizable compound is fed to the reaction zone.This liquid from a distant portion of the reaction medium can becirculated to a location proximate the location of entry of theoxidizable compound via a flow conduit that is disposed internallyand/or externally to the main reaction vessel.

FIGS. 24 and 25 illustrate two preferred methods of circulating liquidfrom a distant portion of the reaction medium to a location near theinlet of the oxidizable compound using an internal (FIG. 24) or external(FIG. 25) conduit. Preferably, the length of the flow conduit from itsinlet (i.e., opening(s) where the liquid enters the conduit) to itsoutlet (i.e., opening(s) where the liquid is discharge from the conduit)is greater than about 1 meter, more preferably greater than about 3meters, still more preferably greater than about 6 meters, and mostpreferably greater than 9 meters. However, the actual length of theconduit becomes less relevant if the liquid is obtained from a separatevessel, perhaps located immediately above or beside the vessel intowhich the oxidizable compound feed is initially released. Liquid fromany separate vessel containing at least some of the reaction medium is apreferred source for initial dilution of the oxidizable compound.

It is preferred that the liquid flowing through the conduit, whateverthe source, has a lower standing concentration of oxidizable compoundthan the reaction medium immediately adjacent to at least one outlet ofthe conduit. Furthermore, it is preferred that the liquid flowingthrough the conduit has a concentration of oxidizable compound in theliquid phase below about 100,000 ppmw, more preferably below about10,000 ppmw, still more preferably below about 1,000 ppmw and mostpreferably below 100 ppmw, where the concentrations are measured beforeaddition to the conduit of the increment of oxidizable compound feed andof any optional, separate solvent feed. When measured after adding theincrement of oxidizable compound feed and optional solvent feed, it ispreferable that the combined liquid stream entering the reaction mediumhas a concentration of oxidizable compound in the liquid phase belowabout 300,000 ppmw, more preferably below about 50,000 ppmw, and mostpreferably below 10,000 ppmw.

It is desirable to maintain the flow through the conduit at a low enoughrate so that the circulated liquid does suppress the desirable overallgradient of oxidizable compound within the reaction medium. In thisregard, it is preferable that the ratio of the mass of the liquid phasein the reaction zone to which the increment of oxidizable compound isinitially released to the mass flow rate of liquid flowing through theconduit be greater than about 0.3 minutes, more preferably greater thanabout 1 minute, still more preferably between about 2 minutes and about120 minutes, and most preferably between 3 minutes and 60 minutes.

There are many means for compelling the liquid to flow through theconduit. Preferred means include gravity, eductors of all typesemploying either gas or liquid or both as the motive fluid, andmechanical pumps of all types. When using an eductor, one embodiment ofthe invention uses as a motive fluid at least one fluid selected fromthe group consisting of: feed of oxidizable compound (liquid or gas),feed of oxidant (gas), feed of solvent (liquid), and a pumped source ofreaction medium (slurry). Another embodiment uses as a motive fluid atleast two fluids selected from the group consisting of: feed ofoxidizable compound, feed of oxidant, and feed of solvent. Still anotherembodiment uses as a motive fluid a combination feed of oxidizablecompound, feed of oxidant, and feed of solvent.

The appropriate diameter or diameters of the circulation conduit mayvary according to the amount and properties of material being conveyed,the energy available for compelling the flow movement, and considerationof capital cost. It is preferable that the minimum diameter for suchconduit is greater than about 0.02 meters, more preferably between about0.06 meters and about 2 meters, and most preferably between 0.12 and 0.8meters As noted above, it is desirable to control flow through theconduit in certain preferred ranges. There are many means known in theart to affect this control by setting an appropriate fixed geometryduring construction of the flow conduit. Another preferred embodiment isto use geometries that are variable during operation, notably includingvalves of all sorts and descriptions, including both manual operationand powered operation by any means, including feed back control loopsfrom a sensing element or without. Another preferred means ofcontrolling the flow of the dilution liquid is to vary the energy inputbetween inlet and outlet of the conduit. Preferred means includechanging the flow rate of one or more motive fluids to an eductor,changing the energy input to a pump driver, and changing the densitydifference or elevation difference when using gravitational force. Thesepreferred means may be used in all combinations as well.

The conduit used for circulation of liquid from the reaction medium maybe of any type known in the art. One embodiment employs a conduitconstructed in whole or part using conventional piping materials.Another embodiment employs a conduit constructed in whole or part usingthe reaction vessel wall as one part of the conduit. A conduit may beconstructed entirely enclosed within the boundaries of the reactionvessel (FIG. 24), or it may be constructed entirely outside the reactionvessel (FIG. 25), or it may comprise sections both within and withoutthe reaction vessel.

The inventors contemplate that, particularly in larger reactors, it maybe desirable to have multiple conduits and of various designs formovement of the liquid through the conduit. Further, it may be desirableto provide multiple outlets at multiple positions on one or all of theconduits. The particulars of the design will balance the desirableoverall gradient in standing concentrations of oxidizable compound withthe desirable initial dilution of oxidizable compound feed, according toother aspects of the current invention.

FIGS. 24 and 25 both illustrate designs that employ a deaeration vesselcoupled to the conduit. This deaeration vessel ensures that the portionof the reaction medium used to dilute the incoming oxidizable compoundis substantially de-aerated slurry. It is now noted, however, that theliquid or slurry used to dilute the incoming oxidizable compound may bein an aerated form as well as a de-aerated form.

The use of a liquid flowing through a conduit to provide dilution of theoxidizable compound feed is particularly useful in bubble columnreactors. Furthermore, in bubble column reactors, a good benefit for theinitial dilution of the oxidizable compound feed can be achieved evenwithout adding the oxidizable compound feed directly into the conduit,providing that the outlet of the conduit lies sufficiently close to theposition of addition of the oxidizable compound. In such an embodiment,it is preferable that the outlet of the conduit be located within about27 conduit outlet diameters of the nearest addition location for theoxidizable compound, more preferably within about 9 conduit outletdiameters, still more preferably within about 3 conduit outletdiameters, and most preferably within 1 conduit outlet diameter.

It has also been discovered that flow eductors can be useful for initialdilution of oxidizable compound feed in oxidation bubble columnsaccording to on embodiment of the present invention, even without theuse of conduits for obtaining dilution liquid from a distant portion ofthe reaction medium. In such cases, the eductor is located within thereaction medium and has an open pathway from the reaction medium intothe throat of the eductor, where low pressure draws in adjacent reactionmedium. Examples of two possible eductor configurations are illustratedin FIGS. 26 and 27. In a preferred embodiment of these eductors, thenearest location of feeding oxidizable compound is within about 4meters, more preferably within about 1 meter and most preferably 0.3meters of the throat of the eductor. In another embodiment, theoxidizable compound is fed under pressure as a motive fluid. In stillanother embodiment, either the solvent or the oxidant is fed underpressure as additional motive fluid along with the oxidizable compound.In yet another embodiment, both the solvent and ant oxidant are fedunder pressure as additional motive fluid along with the oxidizablecompound.

The inventors contemplate that, particularly in larger reactors, it maybe desirable to have multiple eductors and of various designs situatedat various positions within the reaction medium. The particulars of thedesign will balance the desirable overall gradient in standingconcentrations of the oxidizable compound with the desirable initialdilution of the oxidizable compound feed, according to other aspects ofthe current invention. In addition, the inventors contemplate that theoutlet flow plumes from an eductor may be oriented in any direction.When multiple eductors are used, each eductor may be orientedseparately, again in any direction.

As mentioned above, certain physical and operational features of bubblecolumn reactor 20, described above with reference to FIGS. 1-27, providefor vertical gradients in the pressure, temperature, and reactant (i.e.,oxygen and oxidizable compound) concentrations of reaction medium 36. Asdiscussed above, these vertical gradients can provide for a moreeffective and economical oxidation process as compared to conventionaloxidations processes, which favor a well-mixed reaction medium ofrelatively uniform pressure, temperature, and reactant concentrationthroughout. The vertical gradients for oxygen, oxidizable compound(e.g., para-xylene), and temperature made possible by employing anoxidation system in accordance with an embodiment of the presentinvention will now be discussed in greater detail.

Referring now to FIG. 28, in order to quantify the reactantconcentration gradients existing in reaction medium 36 during oxidationin bubble column reactor 20, the entire volume of reaction medium 36 canbe theoretically partitioned into 30 discrete horizontal slices of equalvolume. FIG. 28 illustrates the concept of dividing reaction medium 36into 30 discrete horizontal slices of equal volume. With the exceptionof the highest and lowest horizontal slices, each horizontal slice is adiscrete volume bounded on its top and bottom by imaginary horizontalplanes and bounded on its sides by the wall of reactor 20. The highesthorizontal slice is bounded on its bottom by an imaginary horizontalplane and on its top by the upper surface of reaction medium 36. Thelowest horizontal slice is bounded on its top by an imaginary horizontalplane and on its bottom by the bottom of the vessel shell. Once reactionmedium 36 has been theoretically partitioned into 30 discrete horizontalslices of equal volume, the time-averaged and volume-averagedconcentration of each horizontal slice can then be determined. Theindividual horizontal slice having the maximum concentration of all 30horizontal slices can be identified as the “C-max horizontal slice.” Theindividual horizontal slice located above the C-max horizontal slice andhaving the minimum concentration of all horizontal slices located abovethe C-max horizontal slice can be identified as the “C-min horizontalslice.” The vertical concentration gradient can then be calculated asthe ratio of the concentration in the C-max horizontal slice to theconcentration in the C-min horizontal slice.

With respect to quantifying the oxygen concentration gradient, whenreaction medium 36 is theoretically partitioned into 30 discretehorizontal slices of equal volume, an O₂-max horizontal slice isidentified as having the maximum oxygen concentration of all the 30horizontal slices and an O₂-min horizontal slice is identified as havingthe minimum oxygen concentration of the horizontal slices located abovethe O₂-max horizontal slice. The oxygen concentrations of the horizontalslices are measured in the gas phase of reaction medium 36 on atime-averaged and volume-averaged molar wet basis. It is preferred forthe ratio of the oxygen concentration of the O₂-max horizontal slice tothe oxygen concentration of the O₂-min horizontal slice to be in therange of from about 2:1 to about 25:1, more preferably in the range offrom about 3:1 to about 15:1, and most preferably in the range of from4:1 to 10:1.

Typically, the O₂-max horizontal slice will be located near the bottomof reaction medium 36, while the O₂-min horizontal slice will be locatednear the top of reaction medium 36. Preferably, the O₂-min horizontalslice is one of the 5 upper-most horizontal slices of the 30 discretehorizontal slices. Most preferably, the O₂-min horizontal slice is theupper-most one of the 30 discrete horizontal slices, as illustrated inFIG. 28. Preferably, the O₂-max horizontal slice is one of the 10lower-most horizontal slices of the 30 discrete horizontal slices. Mostpreferably, the O₂-max horizontal slice is one of the 5 lower-mosthorizontal slices of the 30 discrete horizontal slices. For example,FIG. 28 illustrates the O₂-max horizontal slice as the third horizontalslice from the bottom of reactor 20. It is preferred for the verticalspacing between the O₂-min and O₂-max horizontal slices to be at leastabout 2W, more preferably at least about 4W, and most preferably atleast 6W. It is preferred for the vertical spacing between the O₂-minand O₂-max horizontal slices to be at least about 0.2H, more preferablyat least about 0.4H, and most preferably at least 0.6H

The time-averaged and volume-averaged oxygen concentration, on a wetbasis, of the O₂-min horizontal slice is preferably in the range of fromabout 0.1 to about 3 mole percent, more preferably in the range of fromabout 0.3 to about 2 mole percent, and most preferably in the range offrom 0.5 to 1.5 mole percent. The time-averaged and volume-averagedoxygen concentration of the O₂-max horizontal slice is preferably in therange of from about 4 to about 20 mole percent, more preferably in therange of from about 5 to about 15 mole percent, and most preferably inthe range of from 6 to 12 mole percent. The time-averaged concentrationof oxygen, on a dry basis, in the gaseous effluent discharged fromreactor 20 via gas outlet 40 is preferably in the range of from about0.5 to about 9 mole percent, more preferably in the range of from about1 to about 7 mole percent, and most preferably in the range of from 1.5to 5 mole percent.

Because the oxygen concentration decays so markedly toward the top ofreaction medium 36, it is desirable that the demand for oxygen bereduced in the top of reaction medium 36. This reduced demand for oxygennear the top of reaction medium 36 can be accomplished by creating avertical gradient in the concentration of the oxidizable compound (e.g.,para-xylene), where the minimum concentration of oxidizable compound islocated near the top of reaction medium 36.

With respect to quantifying the oxidizable compound (e.g., para-xylene)concentration gradient, when reaction medium 36 is theoreticallypartitioned into 30 discrete horizontal slices of equal volume, anOC-max horizontal slice is identified as having the maximum oxidizablecompound concentration of all the 30 horizontal slices and an OC-minhorizontal slice is identified as having the minimum oxidizable compoundconcentration of the horizontal slices located above the OC-maxhorizontal slice. The oxidizable compound concentrations of thehorizontal slices are measured in the liquid phase on a time-averagedand volume-averaged mass fraction basis. It is preferred for the ratioof the oxidizable compound concentration of the OC-max horizontal sliceto the oxidizable compound concentration of the OC-min horizontal sliceto be greater than about 5:1, more preferably greater than about 10:1,still more preferably greater than about 20:1, and most preferably inthe range of from 40:1 to 1000:1.

Typically, the OC-max horizontal slice will be located near the bottomof reaction medium 36, while the OC-min horizontal slice will be locatednear the top of reaction medium 36. Preferably, the OC-min horizontalslice is one of the 5 upper-most horizontal slices of the 30 discretehorizontal slices. Most preferably, the OC-min horizontal slice is theupper-most one of the 30 discrete horizontal slices, as illustrated inFIG. 28. Preferably, the OC-max horizontal slice is one of the 10lower-most horizontal slices of the 30 discrete horizontal slices. Mostpreferably, the OC-max horizontal slice is one of the 5 lower-mosthorizontal slices of the 30 discrete horizontal slices. For example,FIG. 28 illustrates the OC-max horizontal slice as the fifth horizontalslice from the bottom of reactor 20. It is preferred for the verticalspacing between the OC-min and OC-max horizontal slices to be at leastabout 2W, where “W” is the maximum width of reaction medium 36. Morepreferably, the vertical spacing between the OC-min and OC-maxhorizontal slices is at least about 4W, and most preferably at least 6W.Given a height “H” of reaction medium 36, it is preferred for thevertical spacing between the OC-min and OC-max horizontal slices to beat least about 0.2H, more preferably at least about 0.4H, and mostpreferably at least 0.6H.

The time-averaged and volume-averaged oxidizable compound (e.g.,para-xylene) concentration in the liquid phase of the OC-min horizontalslice is preferably less than about 5,000 ppmw, more preferably lessthan about 2,000 ppmw, still more preferably less than about 400 ppmw,and most preferably in the range of from 1 ppmw to 100 ppmw. Thetime-averaged and volume-averaged oxidizable compound concentration inthe liquid phase of the OC-max horizontal slice is preferably in therange of from about 100 ppmw to about 10,000 ppmw, more preferably inthe range of from about 200 ppmw to about 5,000 ppmw, and mostpreferably in the range of from 500 ppmw to 3,000 ppmw.

Although it is preferred for bubble column reactor 20 to providevertical gradients in the concentration of the oxidizable compound, itis also preferred that the volume percent of reaction medium 36 havingan oxidizable compound concentration in the liquid phase above 1,000ppmw be minimized. Preferably, the time-averaged volume percent ofreaction medium 36 having an oxidizable compound concentration in theliquid phase above 1,000 ppmw is less than about 9 percent, morepreferably less than about 6 percent, and most preferably less than 3percent. Preferably, the time-averaged volume percent of reaction medium36 having an oxidizable compound concentration in the liquid phase above2,500 ppmw is less than about 1.5 percent, more preferably less thanabout 1 percent, and most preferably less than 0.5 percent. Preferably,the time-averaged volume percent of reaction medium 36 having anoxidizable compound concentration in the liquid phase above 10,000 ppmwis less than about 0.3 percent, more preferably less than about 0.1percent, and most preferably less than 0.03 percent. Preferably, thetime-averaged volume percent of reaction medium 36 having an oxidizablecompound concentration in the liquid phase above 25,000 ppmw is lessthan about 0.03 percent, more preferably less than about 0.015 percent,and most preferably less than 0.007 percent. The inventors note that thevolume of reaction medium 36 having the elevated levels of oxidizablecompound need not lie in a single contiguous volume. At many times, thechaotic flow patterns in a bubble column reaction vessel producesimultaneously two or more continuous but segregated portions ofreaction medium 36 having the elevated levels of oxidizable compound. Ateach time used in the time averaging, all such continuous but segregatedvolumes larger than 0.0001 volume percent of the total reaction mediumare added together to determine the total volume having the elevatedlevels of oxidizable compound concentration in the liquid phase.

In addition to the concentration gradients of oxygen and oxidizablecompound, discussed above, it is preferred for a temperature gradient toexist in reaction medium 36. Referring again to FIG. 28, thistemperature gradient can be quantified in a manner similar to theconcentration gradients by theoretically partitioning reaction medium 36into 30 discrete horizontal slices of equal volume and measuring thetime-averaged and volume-averaged temperature of each slice. Thehorizontal slice with the lowest temperature out of the lowest 15horizontal slices can then be identified as the T-min horizontal slice,and the horizontal slice located above the T-min horizontal slice andhaving the maximum temperature of all the slices above the T-minhorizontal slice can then be identified as the “T-max horizontal slice.”It is preferred for the temperature of the T-max horizontal slice be atleast about 1° C. higher than the temperature of the T-min horizontalslice. More preferably the temperature of the T-max horizontal slice isin the range of from about 1.25 to about 12° C. higher than thetemperature of the T-min horizontal slice. Most preferably thetemperature of the T-max horizontal slice is in the range of from 2 to8° C. higher than the temperature of the T-min horizontal slice. Thetemperature of the T-max horizontal slice is preferably in the range offrom about 125 to about 200° C., more preferably in the range of fromabout 140 to about 180° C., and most preferably in the range of from 150to 170° C.

Typically, the T-max horizontal slice will be located near the center ofreaction medium 36, while the T-min horizontal slice will be locatednear the bottom of reaction medium 36. Preferably, the T-min horizontalslice is one of the 10 lower-most horizontal slices of the 15 lowesthorizontal slices. Most preferably, the T-min horizontal slice is one ofthe 5 lower-most horizontal slices of the 15 lowest horizontal slices.For example, FIG. 28 illustrates the T-min horizontal slice as thesecond horizontal slice from the bottom of reactor 20. Preferably, theT-max horizontal slice is one of the 20 middle horizontal slices of the30 discrete horizontal slices. Most preferably, the T-min horizontalslice is one of the 14 middle horizontal slices of the 30 discretehorizontal slices. For example, FIG. 28 illustrates the T-max horizontalslice as the twentieth horizontal slice from the bottom of reactor 20(i.e., one of the middle 10 horizontal slices). It is preferred for thevertical spacing between the T-min and T-max horizontal slices to be atleast about 2W, more preferably at least about 4W, and most preferablyat least 6W. It is preferred for the vertical spacing between the T-minand T-max horizontal slices to be at least about 0.2H, more preferablyat least about 0.4H, and most preferably at least 0.6H.

As discussed above, when a vertical temperature gradient exists inreaction medium 36, it can be advantageous to withdraw reaction medium36 at an elevated location where the temperature of reaction medium ishighest, especially when the withdrawn product is subjected to furtherdownstream processing at higher temperatures. Thus, when reaction medium36 is withdrawn from reaction zone 28 via one or more elevated outlets,as illustrated in FIGS. 19 and 20, it is preferred for the elevatedoutlet(s) to be located near the T-max horizontal slice. Preferably, theelevated outlet is located within 10 horizontal slices of the T-maxhorizontal slice, more preferably within 5 horizontal slices of theT-max horizontal slice, and most preferably within 2 horizontal slicesof the T-max horizontal slice.

It is now noted that many of the inventive features described herein canbe employed in multiple oxidation reactor systems—not just systemsemploying a single oxidation reactor. In addition, certain inventivefeatures described herein can be employed in mechanically-agitatedand/or flow-agitated oxidation reactors—not just bubble-agitatedreactors (i.e., bubble column reactors). For example, the inventors havediscovered certain advantages associated with staging/varying oxygenconcentration and/or oxygen consumption rate throughout the reactionmedium. The advantages realized by the staging of oxygenconcentration/consumption in the reaction medium can be realized whetherthe total volume of the reaction medium is contained in a single vesselor in multiple vessels. Further, the advantages realized by the stagingof oxygen concentration/consumption in the reaction medium can berealized whether the reaction vessel(s) is mechanically-agitated,flow-agitated, and/or bubble-agitated.

One way of quantifying the degree of staging of oxygen concentrationand/or consumption rate in a reaction medium is to compare two or moredistinct 20-percent continuous volumes of the reaction medium. These20-percent continuous volumes need not be defined by any particularshape. However, each 20-percent continuous volume must be formed of acontiguous volume of the reaction medium (i.e., each volume is“continuous”), and the 20-percent continuous volumes must not overlapone another (i.e., the volumes are “distinct”). FIGS. 29-31 illustratethat these distinct 20-percent continuous volumes can be located in thesame reactor (FIG. 29) or in multiple reactors (FIGS. 30 and 31). It isnoted that the reactors illustrated in FIGS. 29-31 can bemechanically-agitated, flow-agitated, and/or bubble-agitated reactors.In one embodiment, it is preferred for the reactors illustrated in FIGS.29-31 to be bubble-agitated reactors (i.e., bubble column reactors).

Referring now to FIG. 29, reactor 20 is illustrated as containing areaction medium 36. Reaction medium 36 includes a first distinct20-percent continuous volume 37 and a second distinct 20-percentcontinuous volume 39.

Referring now to FIG. 30, a multiple reactor system is illustrated asincluding a first reactor 720 a and a second reactor 720 b. Reactors 720a,b cooperatively contain a total volume of a reaction medium 736. Firstreactor 720 a contains a first reaction medium portion 736 a, whilesecond reactor 720 b contains a second reaction medium portion 736 b. Afirst distinct 20-percent continuous volume 737 of reaction medium 736is shown as being defined within first reactor 720 a, while a seconddistinct 20-percent continuous volume 739 of reaction medium 736 isshown as being defined within second reactor 720 b.

Referring now to FIG. 31, a multiple reactor system is illustrated asincluding a first reactor 820 a, a second reactor 820 b, and a thirdreactor 820 c. Reactors 820 a,b,c cooperatively contain a total volumeof a reaction medium 836. First reactor 820 a contains a first reactionmedium portion 836 a; second reactor 820 b contains a second reactionmedium portion 836 b; and third reactor 820 c contains a third reactionmedium portion 836 c. A first distinct 20-percent continuous volume 837of reaction medium 836 is shown as being defined within first reactor820 a; a second distinct 20-percent continuous volume 839 of reactionmedium 836 is shown as being defined within second reactor 820 b; and athird distinct 20-percent continuous volume 841 of reaction medium 836is show as being defined within third reactor 820 c.

The staging of oxygen availability in the reaction medium can bequantified by referring to the 20-percent continuous volume of reactionmedium having the most abundant mole fraction of oxygen in the gas phaseand by referring to the 20-percent continuous volume of reaction mediumhaving the most depleted mole fraction of oxygen in the gas phase. Inthe gas phase of the distinct 20-percent continuous volume of thereaction medium containing the highest concentration of oxygen in thegas phase, the time-averaged and volume-averaged oxygen concentration,on a wet basis, is preferably in the range of from about 3 to about 18mole percent, more preferably in the range of from about 3.5 to about 14mole percent, and most preferably in the range of from 4 to 10 molepercent. In the gas phase of the distinct 20-percent continuous volumeof the reaction medium containing the lowest concentration of oxygen inthe gas phase, the time-averaged and volume-averaged oxygenconcentration, on a wet basis, is preferably in the range of from about0.3 to about 5 mole percent, more preferably in the range of from about0.6 to about 4 mole percent, and most preferably in the range of from0.9 to 3 mole percent. Furthermore, the ratio of the time-averaged andvolume-averaged oxygen concentration, on a wet basis, in the mostabundant 20-percent continuous volume of reaction medium compared to themost depleted 20-percent continuous volume of reaction medium ispreferably in the range of from about 1.5:1 to about 20:1, morepreferably in the range of from about 2:1 to about 12:1, and mostpreferably in the range of from 3:1 to 9:1.

The staging of oxygen consumption rate in the reaction medium can bequantified in terms of an oxygen-STR, initially described above.Oxygen-STR was previously describe in a global sense (i.e., from theperspective of the average oxygen-STR of the entire reaction medium);however, oxygen-STR may also be considered in a local sense (i.e., aportion of the reaction medium) in order to quantify staging of theoxygen consumption rate throughout the reaction medium.

The inventors have discovered that it is very useful to cause theoxygen-STR to vary throughout the reaction medium in general harmonywith the desirable gradients disclosed herein relating to pressure inthe reaction medium and to the mole fraction of molecular oxygen in thegas phase of the reaction medium. Thus, it is preferable that the ratioof the oxygen-STR of a first distinct 20-percent continuous volume ofthe reaction medium compared to the oxygen-STR of a second distinct20-percent continuous volume of the reaction medium be in the range offrom about 1.5:1 to about 20:1, more preferably in the range of fromabout 2:1 to about 12:1, and most preferably in the range of from 3:1 to9:1. In one embodiment the “first distinct 20-percent continuous volume”is located closer than the “second distinct 20-percent continuousvolume” to the location where molecular oxygen is initially introducedinto the reaction medium. These large gradients in oxygen-STR aredesirable whether the partial oxidation reaction medium is contained ina bubble column oxidation reactor or in any other type of reactionvessel in which gradients are created in pressure and/or mole fractionof molecular oxygen in the gas phase of the reaction medium (e.g., in amechanically agitated vessel having multiple, vertically disposedstirring zones achieved by using multiple impellers having strong radialflow, possibly augmented by generally horizontal baffle assemblies, withoxidant flow rising generally upwards from a feed near the lower portionof the reaction vessel, notwithstanding that considerable back-mixing ofoxidant flow may occur within each vertically disposed stirring zone andthat some back-mixing of oxidant flow may occur between adjacentvertically disposed stirring zones). That is, when a gradient exists inthe pressure and/or mole fraction of molecular oxygen in the gas phaseof the reaction medium, the inventors have discovered that it isdesirable to create a similar gradient in the chemical demand fordissolved oxygen by the means disclosed herein.

A preferred means of causing the local oxygen-STR to vary is bycontrolling the locations of feeding the oxidizable compound and bycontrolling the mixing of the liquid phase of the reaction medium tocontrol gradients in concentration of oxidizable compound according toother disclosures of the present invention. Other useful means ofcausing the local oxygen-STR to vary include causing variation inreaction activity by causing local temperature variation and by changingthe local mixture of catalyst and solvent components (e.g., byintroducing an additional gas to cause evaporative cooling in aparticular portion of the reaction medium and by adding a solvent streamcontaining a higher amount of water to decrease activity in a particularportion of the reaction medium).

As discussed above with reference to FIGS. 30 and 31, the partialoxidation reaction can be usefully conducted in multiple reactionvessels wherein at least a portion, preferably at least 25 percent, morepreferably at least 50 percent, and most preferable at least 75 percent,of the molecular oxygen exiting from a first reaction vessel isconducted to one or more subsequent reaction vessels for consumption ofan additional increment, preferably more than 10 percent, morepreferably more than 20 percent, and most preferably more than 40percent, of the molecular oxygen exiting the first/upstream reactionvessel. When using such a series flow of molecular oxygen from onereactor to others, it is desirable that the first reaction vessel isoperated with a higher reaction intensity than at least one of thesubsequent reaction vessels, preferably with the ratio of thevessel-average-oxygen-STR within the first reaction vessel to thevessel-average-oxygen-STR within the subsequent reaction vessel in therange of from about 1.5:1 to about 20:1, more preferably in the range offrom about 2:1 to about 12:1, and most preferably in the range of from3:1 to 9:1.

As discussed above, all types of first reaction vessel (e.g.; bubblecolumn, mechanically-agitated, back-mixed, internally staged, plug flow,and so on) and all types of subsequent reaction vessels, which may ornot be of different type than the first reaction vessel, are useful forseries flow of molecular oxygen to subsequent reaction vessels withaccording to the present invention. The means of causing thevessel-average-oxygen-STR to decline within subsequent reaction vesselsusefully include reduced temperature, reduced concentrations ofoxidizable compound, and reduced reaction activity of the particularmixture of catalytic components and solvent (e.g., reduced concentrationof cobalt, increased concentration of water, and addition of a catalyticretardant such as small quantities of ionic copper).

In flowing from the first reaction vessel to a subsequent reactionvessel, the oxidant stream may be treated by any means known in the artsuch as compression or pressure reduction, cooling or heating, andremoving mass or adding mass of any amount or any type. However, the useof declining vessel-average-oxygen-STR in subsequent reaction vessels isparticularly useful when the absolute pressure in the upper portion ofthe first reaction vessel is less than about 2.0 megapascal, morepreferably less than about 1.6 megapascal, and most preferably less than1.2 megapascal. Furthermore, the use of decliningvessel-average-oxygen-STR in subsequent reaction vessels is particularlyuseful when the ratio of the absolute pressure in the upper portion ofthe first reaction vessel compared to the absolute pressure in the upperportion of at least one subsequent reaction vessel is in the range fromabout 0.5:1 to 6:1, more preferably in a range from about 0.6:1 to about4:1, and most preferably in a range from 0.7:1 to 2:1. Pressurereductions in subsequent vessels below these lower bounds overly reducethe availability of molecular oxygen, and pressure increases above theseupper bounds are increasingly costly compared to using a fresh supply ofoxidant. When using series flow of molecular oxygen to subsequentreaction vessels having declining vessel-average-oxygen-STR, fresh feedstreams of oxidizable compound, solvent and oxidant may flow intosubsequent reaction vessels and/or into the first reaction vessel. Flowsof the liquid phase and the solid phase, if present, of the reactionmedium may flow in any direction between reaction vessels. All or partof the gas phase leaving the first reaction vessel and entering asubsequent reaction vessel may flow separated from or commingled withportions of the liquid phase or the solid phase, if present, of thereaction medium from the first reaction vessel. A flow of product streamcomprising liquid phase and solid phase, if present, may be withdrawnfrom the reaction medium in any reaction vessel in the system.

Referring again to FIGS. 1-31, oxidation is preferably carried out inbubble column reactor 20 under conditions that are markedly different,according to preferred embodiments disclosed herein, than conventionaloxidation reactors. When bubble column reactor 20 is used to carry outthe liquid-phase partial oxidation of para-xylene to crude terephthalicacid (CTA) according to preferred embodiments disclosed herein, thespatial profiles of local reaction intensity, of local evaporationintensity, and of local temperature combined with the liquid flowpatterns within the reaction medium and the preferred, relatively lowoxidation temperatures contribute to the formation of CTA particleshaving unique and advantageous properties.

FIGS. 32A and 32B illustrate base CTA particles produced in accordancewith one embodiment of the present invention. FIG. 32A shows the baseCTA particles at 500 times magnification, while FIG. 32B zooms in on oneof the base CTA particles and shows that particle at 2,000 timesmagnification. As perhaps best illustrated in FIG. 32B, each base CTAparticle is typically formed of a large number of small, agglomeratedCTA subparticles, thereby giving the base CTA particle a relatively highsurface area, high porosity, low density, and good dissolvability. Thebase CTA particles typically have a mean particle size in the range offrom about 20 to about 150 microns, more preferably in the range of fromabout 30 to about 120 microns, and most preferably in the range of from40 to 90 microns. The CTA subparticles typically have a mean particlesize in the range of from about 0.5 to about 30 microns, more preferablyfrom about 1 to about 15 microns, and most preferably in the range offrom 2 to 5 microns. The relatively high surface area of the base CTAparticles illustrated in FIGS. 32A and 32B, can be quantified using aBraunauer-Emmett-Teller (BET) surface area measurement method.Preferably, the base CTA particles have an average BET surface of atleast about 0.6 meters squared per gram (m²/g). More preferably, thebase CTA particles have an average BET surface area in the range of fromabout 0.8 to about 4 m²/g. Most preferably, the base CTA particles havean average BET surface area in the range of from 0.9 to 2 m²/g. Thephysical properties (e.g., particle size, BET surface area, porosity,and dissolvability) of the base CTA particles formed by optimizedoxidation process of a preferred embodiment of the present inventionpermit purification of the CTA particles by more effective and/oreconomical methods, as described in further detail below with respect toFIG. 35.

The mean particle size values provided above were determined usingpolarized light microscopy and image analysis. The equipment employed inthe particle size analysis included a Nikon E800 optical microscope witha 4× Plan Flour N.A. 0.13 objective, a Spot RT™ digital camera, and apersonal computer running Image Pro Plus™ V4.5.0.19 image analysissoftware. The particle size analysis method included the following mainsteps: (1) dispersing the CTA powders in mineral oil; (2) preparing amicroscope slide/cover slip of the dispersion; (3) examining the slideusing polarized light microscopy (crossed polars condition—particlesappear as bright objects on black background); (4) capturing differentimages for each sample preparation (field size=3×2.25 mm; pixelsize=1.84 microns/pixel); (5) performing image analysis with Image ProPlus™ software; (6) exporting the particle measures to a spreadsheet;and (7) performing statistical characterization in the spreadsheet. Step(5) of “performing image analysis with Image Pro Plus™ software”included the substeps of: (a) setting the image threshold to detectwhite particles on dark background; (b) creating a binary image; (c)running a single-pass open filter to filter out pixel noise; (d)measuring all particles in the image; and (e) reporting the meandiameter measured for each particle. The Image Pro Plus™ softwaredefines mean diameter of individual particles as the number averagelength of diameters of a particle measured at 2 degree intervals andpassing through the particle's centroid. Step 7 of “performingstatistical characterization in the spreadsheet” comprises calculatingthe volume-weighted mean particle size as follows. The volume of each ofthe n particles in a sample is calculated as if it were spherical usingpi/6*d_(i)ˆ3; multiplying the volume of each particle times its diameterto find pi/6*d_(i)ˆ4; summing for all particles in the sample of thevalues of pi/6*d_(i){circumflex over ( 4)}; summing the volumes of allparticles in the sample; and calculating the volume-weighted particlediameter as sum for all n particles in the sample of (pi/6*d_(i)ˆ4)divided by sum for all n particles in the sample of (pi/6*d_(i)ˆ3). Asused herein, “mean particle size” refers to the volume-weighted meanparticle size determined according to the above-described test method;and it is also referred to as D(4,3).${D( {4,3} )} = \frac{\sum\limits_{i = 1}^{n}{\frac{\pi}{6}d_{i}^{4}}}{\sum\limits_{i = 1}^{n}{\frac{\pi}{6}d_{i}^{3}}}$

In addition, step 7 comprises finding the particle sizes for whichvarious fractions of the total sample volume are smaller. For example,D(v,0.1) is the particle size for which 10 percent of the total samplevolume is smaller and 90 percent is larger; D(v,0.5) is the particlesize for which one-half of the sample volume is larger and one-half issmaller; D(v,0.9) is the particle size for which 90 percent of the totalsample volume is smaller; and so on. In addition, step 7 comprisescalculating the value of D(v,0.9) minus D(v,0.1), which is hereindefined as the “particle size spread”; and step 7 comprises calculatingthe value of the particle size spread divided by D(4,3), which is hereindefined as the “particle size relative spread.”

Furthermore, it is preferable that the D(v,0.1) of the CTA particles asmeasured above be in the range from about 5 to about 65 microns, morepreferably in the range from about 15 to about 55 microns and mostpreferably in the range from 25 to 45 microns. It is preferable that theD(v,0.5) of the CTA particles as measured above be in the range fromabout 10 to about 90 microns, more preferably in the range from about 20to about 80 microns, and most preferably in the range from 30 to 70microns. It is preferable that the D(v,0.9) of the CTA particles asmeasured above be in the range from about 30 to about 150 microns, morepreferably in the range from about 40 to about 130 microns, and mostpreferably in the range from 50 to 110 microns. It is preferable thatthe particle size relative spread be in the range from about 0.5 toabout 2.0, more preferably in the range from about 0.6 to about 1.5, andmost preferably in the range from 0.7 to 1.3.

The BET surface area values provided above were measured on aMicromeritics ASAP2000 (available from Micromeritics InstrumentCorporation of Norcross, Ga.). In the first step of the measurementprocess, a 2 to 4 gram of sample of the particles was weighed and driedunder vacuum at 50° C. The sample was then placed on the analysis gasmanifold and cooled to 77° K. A nitrogen adsorption isotherm wasmeasured at a minimum of 5 equilibrium pressures by exposing the sampleto known volumes of nitrogen gas and measuring the pressure decline. Theequilibrium pressures were appropriately in the range of P/P₀=0.01-0.20,where P is equilibrium pressure and P₀ is vapor pressure of liquidnitrogen at 77° K. The resulting isotherm was then plotted according tothe following BET equation:$\frac{P}{V_{a}( {P_{o} - P} )} = {\frac{1}{V_{m}C} + {\frac{C - 1}{V_{m}C}( \frac{P}{P_{o}} )}}$where V_(a) is volume of gas adsorbed by sample at P, V_(m) is volume ofgas required to cover the entire surface of the sample with a monolayerof gas, and C is a constant. From this plot, V_(m) and C weredetermined. V_(m) was then converted to a surface area using the crosssectional area of nitrogen at 77° K. by:$A = {\sigma\quad\frac{V_{m}}{RT}}$where σ is cross sectional area of nitrogen at 77° K., T is 77° K. and Ris the gas constant.

As alluded to above, CTA formed in accordance with one embodiment of thepresent invention exhibits superior dissolution properties versesconventional CTA made by other processes. This enhanced dissolution rateallows the inventive CTA to be purified by more efficient and/or moreeffective purification processes. The following description addressesthe manner in which the rate of dissolution of CTA can quantified.

The rate of dissolution of a known amount of solids into a known amountof solvent in an agitated mixture can be measured by various protocols.As used herein, a measurement method called the “timed dissolution test”is defined as follows. An ambient pressure of about 0.1 megapascal isused throughout the timed dissolution test. The ambient temperature usedthroughout the timed dissolution test is about 22° C. Furthermore, thesolids, solvent and all dissolution apparatus are fully equilibratedthermally at this temperature before beginning testing, and there is noappreciable heating or cooling of the beaker or its contents during thedissolution time period. A solvent portion of fresh, HPLC analyticalgrade of tetrahydrofuran (>99.9 percent purity), hereafter THF,measuring 250 grams is placed into a cleaned KIMAX tall form 400milliliter glass beaker (Kimble® part number 14020, Kimble/Kontes,Vineland, N.J.), which is uninsulated, smooth-sided, and generallycylindrical in form. A Teflon-coated magnetic stirring bar (VWR partnumber 58948-230, about 1-inch long with ⅜-inch diameter, octagonalcross section, VWR International, West Chester, Pa. 19380) is placed inthe beaker, where it naturally settles to the bottom. The sample isstirred using a Variomag® multipoint 15 magnetic stirrer (H&PLabortechnik AG, Oberschleissheim, Germany) magnetic stirrer at asetting of 800 revolutions per minute. This stirring begins no more than5 minutes before the addition of solids and continues steadily for atleast 30 minutes after adding the solids. A solid sample of crude orpurified TPA particulates amounting to 250 milligrams is weighed into anon-sticking sample weighing pan. At a starting time designated as t=0,the weighed solids are poured all at once into the stirred THF, and atimer is started simultaneously. Properly done, the THF very rapidlywets the solids and forms a dilute, well-agitated slurry within 5seconds. Subsequently, samples of this mixture are obtained at thefollowing times, measured in minutes from t=0: 0.08, 0.25, 0.50, 0.75,1.00, 1.50, 2.00, 2.50, 3.00, 4.00, 5.00, 6.00, 8.00, 10.00, 15.00, and30.00. Each small sample is withdrawn from the dilute, well-agitatedmixture using a new, disposable syringe (Becton, Dickinson and Co, 5milliliter, REF 30163, Franklin Lakes, N.J. 07417). Immediately uponwithdrawal from the beaker, approximately 2 milliliters of clear liquidsample is rapidly discharged through a new, unused syringe filter (25 mmdiameter, 0.45 micron, Gelman GHP Acrodisc GF®, Pall Corporation, EastHills, N.Y. 11548) into a new, labeled glass sample vial. The durationof each syringe filling, filter placement, and discharging into a samplevial is correctly less than about 5 seconds, and this interval isappropriately started and ended within about 3 seconds either side ofeach target sampling time. Within about five minutes of each filling,the sample vials are capped shut and maintained at approximatelyconstant temperature until performing the following chemical analysis.After the final sample is taken at a time of 30 minutes past t=0, allsixteen samples are analyzed for the amount of dissolved TPA using aHPLC-DAD method generally as described elsewhere within this disclosure.However, in the present test, the calibration standards and the resultsreported are both based upon milligrams of dissolved TPA per gram of THFsolvent (hereafter “ppm in THF”). For example, if all of the 250milligrams of solids were very pure TPA and if this entire amount fullydissolved in the 250 grams of THF solvent before a particular samplewere taken, the correctly measured concentration would be about 1,000ppm in THF.

When CTA according to the present invention is subjected to the timeddissolution test described above, it is preferred that a sample taken atone minute past t=0 dissolves to a concentration of at least about 500ppm in THF, more preferably to at least 600 ppm in THF. For a sampletaken at two minutes past t=0, it is preferred that CTA according to thecurrent invention will dissolve to a concentration of at least about 700ppm in THF, more preferably to at least 750 ppm in THF. For a sampletaken at four minutes past t=0, it is preferred that CTA according tothe current invention will dissolve to a concentration of at least about840 ppm in THF, more preferably to at least 880 ppm in THF.

The inventors have found that a relatively simple negative exponentialgrowth model is useful to describe the time dependence of the entiredata set from a complete timed dissolution test, notwithstanding thecomplexity of the particulate samples and of the dissolution process.The form of the equation, hereinafter the “timed dissolution model”, isas follows:S=A+B*(1−exp(−C*t)), where

-   -   t=time in units of minutes;    -   S=solubility, in units of ppm in THF, at time t;    -   exp=exponential function in the base of the natural logarithm of        2;    -   A, B=regressed constants in units of ppm in THF, where A relates        mostly to the rapid dissolution of the smaller particles at very        short times, and where the sum of A+B relates mostly to the        total amount of dissolution near the end of the specified        testing period; and    -   C=a regressed time constant in units of reciprocal minutes.

The regressed constants are adjusted to minimize the sum of the squaresof the errors between the actual data points and the corresponding modelvalues, which method is commonly called a “least squares” fit. Apreferred software package for executing this data regression is JMPRelease 5.1.2 (SAS Institute Inc., JMP Software, SAS Campus Drive, Cary,N.C. 27513).

When CTA according to the present invention is tested with the timeddissolution test and fitted to the timed dissolution model describedabove, it is preferred for the CTA to have a time constant “C” greaterthan about 0.5 reciprocal minutes, more preferably greater than about0.6 reciprocal minutes, and most preferably greater than 0.7 reciprocalminutes.

FIGS. 33A and 33B illustrate a conventional CTA particle made by aconventional high-temperature oxidation process in a continuous stirredtank reactor (CSTR). FIG. 33A shows the conventional CTA particle at 500times magnification, while FIG. 33B zooms in and shows the CTA particleat 2,000 times magnification. A visual comparison of the inventive CTAparticles illustrated in FIGS. 32A and 32B and the conventional CTAparticle illustrated in FIGS. 33A and 33B shows that the conventionalCTA particle has a higher density, lower surface area, lower porosity,and larger particle size than the inventive CTA particles. In fact, theconventional CTA represented in FIGS. 33A and 33B has a mean particlesize of about 205 microns and a BET surface area of about 0.57 m²/g.

FIG. 34 illustrates a conventional process for making purifiedterephthalic acid (PTA). In the conventional PTA process, para-xylene ispartially oxidized in a mechanically agitated high temperature oxidationreactor 700. A slurry comprising CTA is withdrawn from reactor 700 andthen purified in a purification system 702. The PTA product ofpurification system 702 is introduced into a separation system 706 forseparation and drying of the PTA particles. Purification system 702represents a large portion of the costs associated with producing PTAparticles by conventional methods. Purification system 702 generallyincludes a water addition/exchange system 708, a dissolution system 710,a hydrogenation system 712, and three separate crystallization vessels704 a,b,c. In water addition/exchange system 708, a substantial portionof the mother liquor is displaced with water. After water addition, thewater/CTA slurry is introduced into the dissolution system 710 where thewater/CTA mixture is heated until the CTA particles fully dissolve inthe water. After CTA dissolution, the CTA-in-water solution is subjectedto hydrogenation in hydrogenation system 712. The hydrogenated effluentfrom hydrogenation system 712 is then subjected to three crystallizationsteps in crystallization vessels 704 a,b,c, followed by PTA separationin separation system 706.

FIG. 35 illustrates an improved process for producing PTA employing abubble column oxidation reactor 800 configured in accordance with anembodiment of the present invention. An initial slurry comprising solidCTA particles and a liquid mother liquor is withdrawn from reactor 800.Typically, the initial slurry may contain in the range of from about 10to about 50 weight percent solid CTA particles, with the balance beingliquid mother liquor. The solid CTA particles present in the initialslurry typically contain at least about 400 ppmw of4-carboxybenzaldehyde (4-CBA), more typically at least about 800 ppmw of4-CBA, and most typically in the range of from 1,000 to 15,000 ppmw of4-CBA. The initial slurry withdrawn from reactor 800 is introduced intoa purification system 802 to reduce the concentration of 4-CBA and otherimpurities present in the CTA. A purer/purified slurry is produced frompurification system 802 and is subjected to separation and drying in aseparation system 804 to thereby produce purer solid terephthalic acidparticles comprising less than about 400 ppmw of 4-CBA, more preferablyless than about 250 ppmw of 4-CBA, and most preferably in the range offrom 10 to 200 ppmw of 4-CBA.

Purification system 802 of the PTA production system illustrated in FIG.35 provides a number of advantages over purification system 802 of theprior art system illustrated in FIG. 34. Preferably, purification system802 generally includes a liquor exchange system 806, a digester 808, anda single crystallizer 810. In liquor exchange system 806, at least about50 weight percent of the mother liquor present in the initial slurry isreplaced with a fresh replacement solvent to thereby provide asolvent-exchanged slurry comprising CTA particles and the replacementsolvent. The solvent-exchanged slurry exiting liquor exchange system 806is introduced into digester (or secondary oxidation reactor) 808. Indigester 808, a secondary oxidation reaction is preformed at slightlyhigher temperatures than were used in the initial/primary oxidationreaction carried out in bubble column reactor 800. As discussed above,the high surface area, small particle size, and low density of the CTAparticles produced in reactor 800 cause certain impurities trapped inthe CTA particles to become available for oxidation in digester 808without requiring complete dissolution of the CTA particles in digester808. Thus, the temperature in digester 808 can be lower than manysimilar prior art processes. The secondary oxidation carried out indigester 808 preferably reduces the concentration of 4-CBA in the CTA byat least 200 ppmw, more preferably at least about 400 ppmw, and mostpreferably in the range of from 600 to 6,000 ppmw. Preferably, thesecondary oxidation temperature in digester 808 is at least about 10° C.higher than the primary oxidation temperature in bubble column reactor800, more preferably about 20 to about 80° C. higher than the primaryoxidation temperature in reactor 800, and most preferably 30 to 50° C.higher than the primary oxidation temperature in reactor 800. Thesecondary oxidation temperature is preferably in the range of from about160 to about 240° C., more preferably in the range of from about 180 toabout 220° C. and most preferably in the range of from 190 to 210° C.The purified product from digester 808 requires only a singlecrystallization step in crystallizer 810 prior to separation inseparation system 804. Suitable secondary oxidation/digestion techniquesare discussed in further detail in U.S. Pat. App. Pub. No. 2005/0065373,the entire disclosure of which is expressly incorporated herein byreference.

Terephthalic acid (e.g., PTA) produced by the system illustrated in FIG.35 is preferably formed of PTA particles having a mean particle size ofat least about 40 microns, more preferably in the range of from about 50to about 2,000 microns, and most preferably in the range of from 60 to200 microns. The PTA particles preferably have an average BET surfacearea less than about 0.25 m²/g, more preferably in the range of fromabout 0.005 to about 0.2 m²/g, and most preferably in the range of from0.01 to 0.18 m²/g. PTA produced by the system illustrated in FIG. 35 issuitable for use as a feedstock in the making of PET. Typically, PET ismade via esterification of terephthalic with ethylene glycol, followedby polycondensation. Preferably, terephthalic acid produced by anembodiment of the present invention is employed as a feed to the pipereactor PET process described in U.S. patent application Ser. No.10/013,318, filed Dec. 7, 2001, the entire disclosure of which isincorporated herein by reference.

CTA particles with the preferred morphology disclosed herein areparticularly useful in the above-described oxidative digestion processfor reduction of 4-CBA content. In addition, these preferred CTAparticles provide advantages in a wide range of other post-processesinvolving dissolution and/or chemical reaction of the particles. Theseadditional post-processes include, but are not limited too, reactionwith at least one hydroxyl-containing compound to form ester compounds,especially the reaction of CTA with methanol to form dimethylterephthalate and impurity esters; reaction with at least one diol toform ester monomer and/or polymer compounds, especially the reaction ofCTA with ethylene glycol to form polyethylene terephthalate (PET); andfull or partial dissolution in solvents, including, but not limited too,water, acetic acid, and N-methyl-2-pyrrolidone, which may includefurther processing, including, but not limited too, reprecipitation of amore pure terephthalic acid and/or selective chemical reduction ofcarbonyl groups other than carboxylic acid groups. Notably included isthe substantial dissolution of CTA in a solvent comprising water coupledwith partial hydrogenation that reduces the amount of aldehydes,especially 4-CBA, fluorenones, phenones, and/or anthraquinones.

The inventors also contemplate that CTA particles having the preferredproperties disclosed herein can be produced from CTA particles notconforming to the preferred properties disclosed herein (non-conformingCTA particles) by means including, but not limited too, mechanicalcomminution of non-conforming CTA particles and full or partialdissolution of non-conforming CTA particles followed by full or partialre-precipitation.

In accordance with one embodiment of the present invention, there isprovided a process for partially oxidizing an oxidizable aromaticcompound to one or more types of aromatic carboxylic acid wherein thepurity of the solvent portion of the feed (i.e., the “solvent feed”) andthe purity of the oxidizable compound portion of the feed (i.e., the“oxidizable compound feed”) are controlled within certain rangesspecified below. Along with other embodiments of the present invention,this enables the purity of the liquid phase and, if present, the solidphase and the combined slurry (i.e., solid plus liquid) phase of thereaction medium to be controlled in certain preferred ranges, outlinedbelow.

With respect to the solvent feed, it is known to oxidize an oxidizablearomatic compound(s) to produce an aromatic carboxylic acid wherein thesolvent feed introduced into the reaction medium is a mixture ofanalytical-purity acetic acid and water, as is often employed atlaboratory scale and pilot scale. Likewise, it is known to conduct theoxidation of oxidizable aromatic compound to aromatic carboxylic acidwherein the solvent leaving the reaction medium is separated from theproduced aromatic carboxylic acid and then recycled back to the reactionmedium as feed solvent, primarily for reasons of manufacturing cost.This solvent recycling causes certain feed impurities and processby-products to accumulate over time in the recycled solvent. Variousmeans are known in the art to help purify recycled solvent beforere-introduction into the reaction medium. Generally, a higher degree ofpurification of the recycled solvent leads to significantly highermanufacturing cost than does a lower degree of purification by similarmeans. One embodiment of the present invention relates to understandingand defining the preferred ranges of a large number of impurities withinthe solvent feed, many of which were heretofore thought largely benign,in order to find an optimal balance between overall manufacturing costand overall product purity. “Recycled solvent feed” is defined herein assolvent feed comprising at least about 5 weight percent mass that haspreviously passed through a reaction medium containing one or moreoxidizable aromatic compounds undergoing partial oxidation. For reasonsof solvent inventory and of on-stream time in a manufacturing unit, itis preferable that portions of recycled solvent pass through reactionmedium at least once per day of operation, more preferably at least onceper day for at least seven consecutive days of operation, and mostpreferably at least once per day for at least 30 consecutive days ofoperation. For economic reasons, it is preferable that at least about 20weight percent of the solvent feed to the reaction medium of the presentinvention is recycled solvent, more preferably at least about 40 weightpercent, still more preferably at least about 80 weight percent, andmost preferably at least 90 weight percent.

The inventors have discovered that, for reasons of reaction activity andfor consideration of metallic impurities left in the oxidation product,the concentrations of selected multivalent metals within the recycledsolvent feed are preferably in ranges specified immediately below. Theconcentration of iron in recycled solvent is preferably below about 150ppmw, more preferably below about 40 ppmw, and most preferably between 0and 8 ppmw. The concentration of nickel in recycled solvent ispreferably below about 150 ppmw, more preferably below about 40 ppmw,and most preferably between 0 and 8 ppmw. The concentration of chromiumin recycled solvent is preferably below about 150 ppmw, more preferablybelow about 40 ppmw, and most preferably between 0 and 8 ppmw. Theconcentration of molybdenum in recycled solvent is preferably belowabout 75 ppmw, more preferably below about 20 ppmw, and most preferablybetween 0 and 4 ppmw. The concentration of titanium in recycled solventis preferably below about 75 ppmw, more preferably below about 20 ppmw,and most preferably between 0 and 4 ppmw. The concentration of copper inrecycled solvent is preferably below about 20 ppmw, more preferablybelow about 4 ppmw, and most preferably between 0 and 1 ppmw. Othermetallic impurities are also typically present in recycled solvent,generally varying at lower levels in proportion to one or more of theabove listed metals. Controlling the above listed metals in thepreferred ranges will keep other metallic impurities at suitable levels.

These metals can arise as impurities in any of the incoming processfeeds (e.g., in incoming oxidizable compound, solvent, oxidant, andcatalyst compounds). Alternatively, the metals can arise as corrosionproducts from any of the process units contacting reaction medium and/orcontacting recycled solvent. The means for controlling the metals in thedisclosed concentration ranges include the appropriate specification andmonitoring of the purity of various feeds and the appropriate usage ofmaterials of construction, including, but not limited to, manycommercial grades of titanium and of stainless steels including thosegrades known as duplex stainless steels and high molybdenum stainlesssteels.

The inventors have also discovered preferred ranges for selectedaromatic compounds in the recycled solvent. These include bothprecipitated and dissolved aromatic compounds within the recycledsolvent.

Surprisingly, even precipitated product (e.g., TPA) from a partialoxidation of para-xylene, is a contaminant to be managed in recycledsolvent. Because there are surprisingly preferred ranges for the levelsof solids within the reaction medium, any precipitated product in thesolvent feed directly subtracts from the amount of oxidizable compoundthat can be fed in concert. Furthermore, feeding precipitated TPA solidsin the recycled solvent at elevated levels has been discovered to affectadversely the character of the particles formed within a precipitatingoxidation medium, leading to undesirable character in downstreamoperations (e.g., product filtration, solvent washing, oxidativedigestion of crude product, dissolution of crude product for furtherprocessing, and so on). Another undesirable characteristic ofprecipitated solids in the recycle solvent feed is that these oftencontain very high levels of precipitated impurities, as compared toimpurity concentrations in the bulk of the solids within the TPAslurries from which much of the recycled solvent is obtained. Possibly,the elevated levels of impurities observed in solids suspended inrecycled filtrate may relate to nucleation times for precipitation ofcertain impurities from the recycled solvent and/or to cooling of therecycled solvent, whether intentional or due to ambient losses. Forexample, concentrations of highly-colored and undesirable2,6-dicarboxyfluorenone have been observed at far higher levels insolids present in recycled solvent at 80° C. than are observed in TPAsolids separated from recycled solvent at 160° C. Similarly,concentrations of isophthalic acid have been observed at much higherlevels in solids present in recycled solvent compared to levels observedin TPA solids from the reaction medium. Exactly how specificprecipitated impurities entrained within recycled solvent behave whenre-introduced to the reaction medium appears to vary. This dependsperhaps upon the relative solubility of the impurity within the liquidphase of the reaction medium, perhaps upon how the precipitated impurityis layered within the precipitated solids, and perhaps upon the localrate of TPA precipitation where the solid first re-enters the reactionmedium. Thus, the inventors have found it useful to control the level ofcertain impurities in the recycled solvent, as disclosed below, withoutrespect to whether these impurities are present in the recycled solventin dissolved form or are entrained particulates therein.

The amount of precipitated solids present in recycled filtrate isdetermined by a gravimetric method as follows. A representative sampleis withdrawn from the solvent supply to the reaction medium while thesolvent is flowing in a conduit toward the reaction medium. A usefulsample size is about 100 grams captured in a glass container havingabout 250 milliliters of internal volume. Before being released toatmospheric pressure, but while continuously flowing toward the samplecontainer, the recycled filtrate is cooled to less than 100° C.; thiscooling is in order to limit solvent evaporation during the shortinterval before being sealed closed in the glass container. After thesample is captured at atmospheric pressure, the glass container issealed closed immediately. Then the sample is allowed to cool to about20° C. while surrounded by air at about 20° C. and without forcedconvection. After reaching about 20° C., the sample is held at thiscondition for at least about 2 hours. Then, the sealed container isshaken vigorously until a visibly uniform distribution of solids isobtained. Immediately thereafter, a magnetic stirrer bar is added to thesample container and rotated at sufficient speed to maintain effectivelyuniform distribution of solids. A 10 milliliter aliquot of the mixedliquid with suspended solids is withdrawn by pipette and weighed. Thenthe bulk of the liquid phase from this aliquot is separated by vacuumfiltration, still at about 20° C. and effectively without loss ofsolids. The moist solids filtered from this aliquot are then dried,effectively without sublimation of solids, and these dried solids areweighed. The ratio of the weight of the dried solids to the weight ofthe original aliquot of slurry is the fraction of solids, typicallyexpressed as a percentage and referred to herein as the recycledfiltrate content of precipitated solids at 20° C.

The inventors have discovered that aromatic compounds dissolved in theliquid phase of the reaction medium and comprising aromatic carboxylicacids lacking non-aromatic hydrocarbyl groups (e.g., isophthalic acid,benzoic acid, phthalic acid, 2,5,4′-tricarboxybiphenyl) are surprisinglypernicious components. Although these compounds are much reduced inchemical activity in the subject reaction medium compared to oxidizablecompounds having non-aromatic hydrocarbyl groups, the inventors havediscovered that these compounds nonetheless undergo numerous detrimentalreactions. Thus, it is advantageous to control the content of thesecompounds in preferred ranges in the liquid phase of the reactionmedium. This leads to preferred ranges of select compounds in recycledsolvent feed and also to preferred ranges of select precursors in theoxidizable aromatic compound feed.

For example, in the liquid-phase partial oxidation of para-xylene toterephthalic acid (TPA), the inventors have discovered that thehighly-colored and undesirable impurity 2,7-dicarboxyfluorenone(2,7-DCF) is virtually undetectable in the reaction medium and productoff-take when meta-substituted aromatic compounds are at very low levelsin the reaction medium. The inventors have discovered that whenisophthalic acid impurity is present at increasing levels in the solventfeed, the formation of 2,7-DCF rises in almost direct proportion. Theinventors have also discovered that when meta-xylene impurity is presentin the feed of para-xylene, the formation of 2,7-DCF again rises almostin direct proportion. Furthermore, even if the solvent feed andoxidizable compound feed are devoid of meta-substituted aromaticcompounds, the inventors have discovered that some isophthalic acid isformed during a typical partial oxidation of very pure para-xylene,particularly when benzoic acid is present in the liquid phase of thereaction medium. This self-generated isophthalic acid may, owing to itsgreater solubility than TPA in solvent comprising acetic acid and water,build up over time in commercial units employing recycled solvent. Thus,the amount of isophthalic acid within solvent feed, the amount ofmeta-xylene within oxidizable aromatic compound feed, and the rate ofself-creation of isophthalic acid within the reaction medium are allappropriately considered in balance with each other and in balance withany reactions that consume isophthalic acid. Isophthalic acid has beendiscovered to undergo additional consumptive reactions besides theformation of 2,7-DCF, as are disclosed below. In addition, the inventorshave discovered that there are other issues to consider when settingappropriate ranges for the meta-substituted aromatic species in thepartial oxidation of para-xylene to TPA. Other highly-colored andundesirable impurities, such as 2,6-dicarboxyfluorenone (2,6-DCF),appear to relate greatly to dissolved, para-substituted aromaticspecies, which are always present with para-xylene feed to aliquid-phase oxidation. Thus, the suppression of 2,7-DCF is bestconsidered in perspective with the level of other colored impuritiesbeing produced.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that the formation of trimelliticacid rises as the levels isophthalic acid and phthalic acid rise withinthe reaction medium. Trimellitic acid is a tri-functional carboxylicacid leading to branching of polymer chains during production of PETfrom TPA. In many PET applications, branching levels must be controlledto low levels and hence trimellitic acid must be controlled to lowlevels in purified TPA. Besides leading to trimellitic acid, thepresence of meta-substituted and ortho-substituted species in thereaction medium also give rise to other tricarboxylic acids (e.g.,1,3,5-tricarboxybenzene). Furthermore, the increased presence oftricarboxylic acids in the reaction medium increases the amount oftetracarboxylic acid formation (e.g., 1,2,4,5-tetracarboxybenzene).Controlling the summed production of all aromatic carboxylic acidshaving more than two carboxylic acid groups is one factor in setting thepreferred levels of meta-substituted and ortho-substituted species inthe recycled solvent feed, in the oxidizable compound feed, and in thereaction medium according to the present invention.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that increased levels in the liquidphase of the reaction medium of several dissolved aromatic carboxylicacids lacking non-aromatic hydrocarbyl groups leads directly to theincreased production of carbon monoxide and carbon dioxide. Thisincreased production of carbon oxides represents a yield loss on bothoxidant and on oxidizable compound, the later since many of theco-produced aromatic carboxylic acids, which on the one hand may beviewed as impurities, on the other hand also have commercial value.Thus, appropriate removal of relatively soluble carboxylic acids lackingnon-aromatic hydrocarbyl groups from recycle solvent has an economicvalue in preventing yield loss of oxidizable aromatic compound and ofoxidant, in addition to suppressing the generation of highly undesirableimpurities such as various fluorenones and trimellitic acid.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that formation of2,5,4′-tricarboxybiphenyl is seemingly unavoidable. The2,5,4′-tricarboxybiphenyl is an aromatic tricarboxylic acid formed bythe coupling of two aromatic rings, perhaps by the coupling of adissolved para-substituted aromatic species with an aryl radical,perhaps an aryl radical formed by decarboxylation or decarbonylation ofa para-substituted aromatic species. Fortunately, the2,5,4′-tricarboxybiphenyl is typically produced at lower levels thantrimellitic acid and does not usually lead to significantly increaseddifficulties with branching of polymer molecules during production ofPET. However, the inventors have discovered that elevated levels of2,5,4′-tricarboxybiphenyl in a reaction medium comprising oxidation ofalkyl aromatics according to preferred embodiments of the presentinvention lead to increased levels of highly-colored and undesirable2,6-DCF. The increased 2,6-DCF is possibly created from the2,5,4′-tricarboxybiphenyl by ring closure with loss of a water molecule,though the exact reaction mechanism is not known with certainty. If2,5,4′-tricarboxybiphenyl, which is more soluble in solvent comprisingacetic acid and water than is TPA, is allowed to build up too highwithin recycled solvent, conversion rates to 2,6-DCF can becomeunacceptably large.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that aromatic carboxylic acidslacking non-aromatic hydrocarbyl groups (e.g., isophthalic acid)generally lead to mild suppression of the chemical activity of thereaction medium when present in the liquid phase at sufficientconcentration.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that precipitation is very oftennon-ideal (i.e. non-equilibrium) with respect to the relativeconcentrations of different chemical species in the solid phase and inthe liquid phase. Perhaps, this is because the precipitation rate isvery fast at the space-time reaction rates preferred herein, leading tonon-ideal co-precipitation of impurities, or even occlusion. Thus, whenit is desired to limit the concentration of certain impurities (e.g.,trimellitic acid and 2,6-DCF) within crude TPA, owing to theconfiguration of downstream unit operations, it is preferable to controltheir concentration in solvent feed as well as their generation ratewithin the reaction medium.

For example, the inventors have discovered that benzophenone compounds(e.g., 4,4′-dicarboxybenzophenone and 2,5,4′-tricarboxybenzophenone)made during partial oxidation of para-xylene, have undesirable effectsin a PET reaction medium even though benzophenone compounds are not ashighly colored in TPA per se as are fluorenones and anthraquinones.Accordingly, it is desirable to limit the presence of benzophenones andselect precursors in recycled solvent and in oxidizable compound feed.Furthermore, the inventors have discovered that the presence of elevatedlevels of benzoic acid, whether admitted in recycled solvent or formedwithin the reaction medium, leads to elevated rates of production of4,4′-dicarboxybenzophenone.

In review, the inventors have discovered and sufficiently quantified asurprising array of reactions for aromatic compounds lackingnon-aromatic hydrocarbyl groups that are present in the liquid-phasepartial oxidation of para-xylene to TPA. Recapping just the single caseof benzoic acid, the inventors have discovered that increased levels ofbenzoic acid in the reaction medium of certain embodiments of thepresent invention lead to greatly increased production of the highlycolored and undesirable 9-fluorenone-2-carboxylic acid, to greatlyincreased levels of 4,4′-dicarboxybiphenyl, to increased levels of4,4′-dicarboxybenzophenone, to a mild suppression of chemical activityof the intended oxidation of para-xylene, and to increased levels ofcarbon oxides and attendant yield losses. The inventors have discoveredthat increased levels of benzoic acid in the reaction medium also leadto increased production of isophthalic acid and phthalic acid, thelevels of which are desirably controlled in low ranges according tosimilar aspects of the current invention. The number and importance ofreactions involving benzoic acid are perhaps even more surprising sincesome recent inventors contemplate using benzoic acid in place of aceticacid as a primary component of solvent (See, e.g., U.S. Pat. No.6,562,997). Additionally, the present inventors have observed thatbenzoic acid is self-generated during oxidation of para-xylene at ratesthat are quite important relative to its formation from impurities, suchas toluene and ethylbenzene, commonly found in oxidizable compound feedcomprising commercial-purity para-xylene.

On the other hand, the inventors have discovered little value fromadditional regulation of recycled solvent composition in regard to thepresence of oxidizable aromatic compound and in regard to aromaticreaction intermediates that both retain non-aromatic hydrocarbyl groupsand are also relatively soluble in the recycled solvent. In general,these compounds are either fed to or created within the reaction mediumat rates substantially greater than their presence in recycled solvent;and the consumption rate of these compounds within the reaction mediumis great enough, retaining one or more non-aromatic hydrocarbyl groups,to limit appropriately their build-up within recycled solvent. Forexample, during partial oxidation of para-xylene in a multi-phasereaction medium, para-xylene evaporates to a limited extent along withlarge quantities of solvent. When this evaporated solvent exits thereactor as part of the off-gas and is condensed for recovery as recycledsolvent, a substantial portion of the evaporated para-xylene condensestherein as well. It is not necessary to limit the concentration of thispara-xylene in recycled solvent. For example, if solvent is separatedfrom solids upon slurry exiting a para-xylene oxidation reaction medium,this recovered solvent will contain a similar concentration of dissolvedpara-toluic acid to that present at the point of removal from thereaction medium. Although it may be important to limit the standingconcentration of para-toluic acid within the liquid phase of thereaction medium, see below, it is not necessary to regulate separatelythe para-toluic acid in this portion of recycled solvent owing to itsrelatively good solubility and to its low mass flow rate relative to thecreation of para-toluic acid within the reaction medium. Similarly, theinventors have discovered little reason to limit the concentrations inrecycled solvent of aromatic compounds with methyl substituents (e.g.toluic acids), aromatic aldehydes (e.g., terephthaldehyde), of aromaticcompounds with hydroxy-methyl substituents (e.g., 4-hydroxymethylbenzoicacid), and of brominated aromatic compounds retaining at least onenon-aromatic hydrocarbyl group (e.g., alpha-bromo-para-toluic acid)below those inherently found in the liquid phase exiting from thereaction medium occurring in the partial oxidation of xylene accordingto preferred embodiments of the present invention. Surprisingly, theinventors have also discovered that it is also not necessary to regulatein recycled solvent the concentration of selected phenols intrinsicallyproduced during partial oxidation of xylene, for these compounds arecreated and destroyed within the reaction medium at rates much greaterthan their presence in recycled solvent. For example, the inventors havediscovered that 4-hydroxybenzoic acid has relatively small effects onchemical activity in the preferred embodiments of the present inventionwhen co-fed at rates of over 2 grams of 4-hydroxybenzoic acid per 1kilogram of para-xylene, far higher than the natural presence inrecycled solvent, despite being reported by others as a significantpoison in similar reaction medium (See, e.g., W. Partenheimer, CatalysisToday 23 (1995) p. 81).

Thus, there are numerous reactions and numerous considerations insetting the preferred ranges of various aromatic impurities in thesolvent feed as now disclosed. These discoveries are stated in terms ofthe aggregated weight average composition of all solvent streams beingfed to the reaction medium during the course of a set time period,preferably one day, more preferably one hour, and most preferably oneminute. For example, if one solvent feed flows substantiallycontinuously with a composition of 40 ppmw of isophthalic acid at a flowrate of 7 kilograms per minute, a second solvent feed flowssubstantially continuously with a composition of 2,000 ppmw ofisophthalic acid at a flow rate of 10 kilograms per minute, and thereare no other solvent feed streams entering the reaction medium, then theaggregated weight average composition of the solvent feed is calculatedas (40*7+2,000*10)/(7+10)=1,193 ppmw of isophthalic acid. It is notablethat the weight of any oxidizable compound feed or of any oxidant feedthat are perhaps commingled with the solvent feed before entering thereaction medium are not considered in calculating the aggregated weightaverage composition of the solvent feed.

Table 1, below, lists preferred values for certain components in thesolvent feed introduced into the reaction medium. The solvent feedcomponents listed in Table 1 are as follows: 4-carboxybenzaldehyde(4-CBA), 4,4′-dicarboxystilbene (4,4′-DCS), 2,6-dicarboxyanthraquinone(2,6-DCA), 2,6-dicarboxyfluorenone (2,6-DCF), 2,7-dicarboxyfluorenone(2,7-DCF), 3,5-dicarboxyfluorenone (3,5-DCF), 9-fluorenone-2-carboxylicacid (9F-2CA), 9-fluorenone-4-carboxylic acid (9F-4CA), totalfluorenones including other fluorenones not individually listed (totalfluorenones), 4,4′-dicarboxybiphenyl (4,4′-DCB),2,5,4′-tricarboxybiphenyl (2,5,4′-TCB), phthalic acid (PA), isophthalicacid (IPA), benzoic acid (BA), trimellitic acid (TMA),2,6-dicarboxybenzocoumarin (2,6-DCBC), 4,4′-dicarboxybenzil (4,4′-DCBZ),4,4′-dicarboxybenzophenone (4,4′-DCBP), 2,5,4′-tricarboxybenzophenone(2,5,4′-TCBP), terephthalic acid (TPA), precipitated solids at 20° C.,and total aromatic carboxylic acids lacking non-aromatic hydrocarbylgroups. Table 1, below provides the preferred amounts of theseimpurities in CTA produced according to an embodiment of the presentinvention. TABLE 1 Components of Solvent Feed Introduced into ReactionMedium Component Preferred More Preferred Most Preferred IdentificationAmt. (ppmw) Amt. (ppmw) Amt. (ppmw) 4-CBA <1,200  30-600  60-3004,4′-DCS <3 <2 <1 2,6-DCA <6 0.1-3   0.2-1   2,6-DCF <20 0.1-10  0.5-5  2,7-DCF <10 0.1-5   0.5-2   3,5-DCF <10 <5 <2 9F-2CA <10 0.1-5   0.5-2  9F-4CA <5 <3 <1 Total fluorenones <40 <20 1-8 4,4′-DCB <45 <15 0.5-5  2,5,4′-TCB <45 0.1-15  0.5-5   PA <1,000  15-400  40-150 IPA 2,500  40-1,200 120-400 BA <4,500   50-1,500 150-500 TMA <1,000  15-400 40-150 2,6-DCBC <40 <20 <5 4,4′-DCBZ <40 <20 <5 4,4′-DCBP <40 <20 <52,5,4′-TCBP <40 <20 0.5-5   TPA <9,000   200-6,000   400-2,000Precipitated <9,000   200-6,000   600-2,000 Solids at 20° C. TotalAromatic <18,000   300-9,000   450-3,000 Carboxylic Acids Lacking Non-Aromatic Hydrocarbyl Groups

Many other aromatic impurities are also typically present in recycledsolvent, generally varying at even lower levels and/or in proportion toone or more of the disclosed aromatic compounds. Methods for controllingthe disclosed aromatic compounds in the preferred ranges will typicallykeep other aromatic impurities at suitable levels.

When bromine is used within the reaction medium, a large number of ionicand organic forms of bromine are known to exist in a dynamicequilibrium. These various forms of bromine have different stabilitycharacteristics once leaving the reaction medium and passing throughvarious unit operations pertaining to recycled solvent. For example,alpha-bromo-para-toluic acid may persist as such at some conditions ormay rapidly hydrolyze at other conditions to form 4-hydroxymethylbenzoicacid and hydrogen bromide. In the present invention, it is preferablethat at least about 40 weight percent, more preferable that at leastabout 60 weight percent, and most preferable that at least about 80weight percent of the total mass of bromine present in the aggregatedsolvent feed to the reaction medium is in one or more of the followingchemical forms: ionic bromine, alpha-bromo-para-toluic acid, andbromoacetic acid.

Although the importance and value of controlling the aggregated weightaverage purity of solvent feed within the disclosed, desired ranges ofthe present invention has not heretofore been discovered and/ordisclosed, suitable means for controlling the solvent feed purity may beassembled from various methods already known in the art. First, anysolvent evaporated from the reaction medium is typically of suitablepurity providing that liquid or solids from the reaction medium are notentrained with the evaporated solvent. The feeding of reflux solventdroplets into the off-gas disengaging space above the reaction medium,as disclosed herein, appropriately limits such entrainment; and recycledsolvent of suitable purity with respect to aromatic compound can becondensed from such off-gas. Second, the more difficult and costlypurification of recycled solvent feed typically relates to solvent takenfrom the reaction medium in liquid form and to solvent that subsequentlycontacts the liquid and/or solid phases of the reaction medium withdrawnfrom the reaction vessel (e.g., recycled solvent obtained from a filterin which solids are concentrated and/or washed, recycled solventobtained from a centrifuge in which solids are concentrated and/orwashed, recycled solvent taken from a crystallization operation, and soon). However, means are also known in the art for effecting thenecessary purification of these recycled solvent streams using one ormore prior disclosures. With respect to controlling precipitated solidsin recycled solvent to be within the ranges specified, suitable controlmeans include, but are not limited to, gravimetric sedimentation,mechanical filtration using filter cloth on rotary belt filters androtary drum filters, mechanical filtration using stationary filtermedium within pressure vessels, hydro-cyclones, and centrifuges. Withrespect to controlling dissolved aromatic species in recycled solvent tobe within the ranges specified, the control means include, but are notlimited to, those disclosed in U.S. Pat. No. 4,939,297 and U.S. Pat.App. Pub. No. 2005-0038288, incorporated herein by reference. However,none of these prior inventions discovered and disclosed the preferredlevels of purity in the aggregated solvent feed as disclosed herein.Rather, these prior inventions merely provided means to purify selectedand partial streams of recycled solvent without deducing the presentinventive, optimal values of the composition of the aggregated weightaverage solvent feed to the reaction medium.

Turning now to the purity of the feed of oxidizable compound, it isknown that certain levels of isophthalic acid, phthalic acid, andbenzoic acid are present and tolerable at low levels in purified TPAused for polymer production. Moreover, it is known these species arerelatively more soluble in many solvents and may be advantageouslyremoved from purified TPA by crystallization processes. However, from anembodiment of the invention disclosed herein, it is now known thatcontrolling the level of several relatively soluble aromatic species,notably including isophthalic acid, phthalic acid, and benzoic acid, inthe liquid phase of the reaction medium is surprisingly important forcontrolling the level of polycyclic and colored aromatic compoundscreated in the reaction medium, for controlling compounds with more than2 carboxylic acid functions per molecule, for controlling reactionactivity within the partial oxidation reaction medium, and forcontrolling yield losses of oxidant and of aromatic compound.

It is known within the art that isophthalic acid, phthalic acid, andbenzoic acid are formed in the reaction medium as follows. Meta-Xylenefeed impurity oxidizes in good conversion and yield to IPA. Ortho-Xylenefeed impurity oxidizes in good conversion and yield to phthalic acid.Ethylbenzene and toluene feed impurities oxidize in good conversion andyield to benzoic acid. However, the inventors have observed thatsignificant amounts of isophthalic acid, phthalic acid, and benzoic acidare also formed within a reaction medium comprising para-xylene by meansother than oxidation of meta-xylene, ortho-xylene, ethylbenzene, andtoluene. These other intrinsic chemical routes possibly includedecarbonylation, decarboxylation, the re-organization of transitionstates, and addition of methyl and carbonyl radicals to aromatic rings.

In determining preferred ranges of impurities in the feed of oxidizablecompound, many factors are relevant. Any impurity in the feed is likelyto be a direct yield loss and a product purification cost if the purityrequirements of the oxidized product are sufficiently strict (e.g., in areaction medium for partial oxidation of para-xylene, toluene andethylbenzene typically found in commercial-purity para-xylene lead tobenzoic acid, and this benzoic acid is largely removed from mostcommercial TPA). When the partial oxidation product of a feed impurityparticipates in additional reactions, factors other than simple yieldloss and removal become appropriate when considering how much feedpurification cost to incur (e.g., in a reaction medium for partialoxidation of para-xylene, ethylbenzene leads to benzoic acid, andbenzoic acid subsequently leads to highly colored9-fluorenone-2-carboxylic acid, to isophthalic acid, to phthalic acid,and to increased carbon oxides, among others). When the reaction mediumself-generates additional amounts of an impurity by chemical mechanismsnot directly related to feed impurities, the analysis becomes still morecomplex (e.g., in a reaction medium for partial oxidation ofpara-xylene, benzoic acid is also self-generated from para-xyleneitself). In addition, the downstream processing of the crude oxidationproduct may affect the considerations for preferred feed purity. Forexample, the cost of removing to suitable levels a direct impurity(benzoic acid) and subsequent impurities (isophthalic acid, phthalicacid, 9-fluorenone-2-carboxylic acid, et al.) may be one and the same,may be different from each other, and may be different from therequirements of removing a largely unrelated impurity (e.g., incompleteoxidation product 4-CBA in the oxidation of para-xylene to TPA).

The following disclosed feed purity ranges for para-xylene are preferredwhere para-xylene is fed with solvent and oxidant to a reaction mediumfor partial oxidation to produce TPA. These ranges are more preferred inTPA production process having post-oxidation steps to remove fromreaction medium impurities other than oxidant and solvent (e.g.,catalyst metals). These ranges are still more preferred in TPAproduction processes that remove additional 4-CBA from CTA (e.g., byconversion of CTA to dimethyl terephthalate plus impurity esters andsubsequent separation of the methyl ester of 4-CBA by distillation, byoxidative digestion methods for converting 4-CBA to TPA, byhydrogenation methods for converting 4-CBA to para-toluic acid, which isthen separated by partial-crystallization methods). These ranges aremost preferred in TPA production processes that remove additional 4-CBAfrom CTA by oxidative digestion methods for converting 4-CBA to TPA.

Using new knowledge of preferred ranges of recycling aromatic compoundsand of the relative amounts of the aromatic compounds formed directlyfrom oxidation of feed impurities as compared to other intrinsicchemical routes, improved ranges for impurities have been discovered forimpure para-xylene being fed to a partial oxidation process for TPAproduction. Table 2, below provides preferred values for the amount ofmeta-xylene, ortho-xylene, and ethylbenzene+toluene in the para-xylenefeed. TABLE 2 Components of Impure para-xylene Feed Preferred ComponentAmt. More Preferred Most Preferred Identification (ppmw) Amt. (ppmw)Amt. (ppmw) meta-xylene 20-800 50-600 100-400 ortho-xylene 10-300 20-200 30-100 ethylbenzene + toluene* 20-700 50-500 100-300 total 50-900100-800  200-700*Specification for ethylbenzene + toluene is each separately and in sum

Those skilled in the art will now recognize the above impurities withinimpure para-xylene may have their greatest effect on the reaction mediumafter their partial oxidation products have accumulated in recycledsolvent. For example, feeding the upper amount of the most preferredrange of meta-xylene, 400 ppmw, will immediately produce about 200 ppmwof isophthalic acid within the liquid phase of the reaction medium whenoperating with about 33 weight percent solids in the reaction medium.This compares with an input from the upper amount of the most preferredrange for isophthalic acid in recycled solvent of 400 ppmw which, afterallowing for a typical solvent evaporation to cool the reaction medium,amounts to about 1,200 ppmw of isophthalic acid within the liquid phaseof the reaction medium. Thus, it is the accumulation of partialoxidation products over time within recycled solvent that represents thegreatest probable impact of the meta-xylene, ortho-xylene, ethylbenzene,and toluene impurities in the feed of impure para-xylene. Accordingly,the above ranges for impurities in impure para-xylene feed are preferredto be maintained for at least one-half of each day of operation of anypartial oxidation reaction medium in a particular manufacturing unit,more preferably for at least three-quarters of each day for at leastseven consecutive days of operation, and most preferably when themass-weighted averages of the impure para-xylene feed composition arewithin the preferred ranges for at least 30 consecutive days ofoperation.

Means for obtaining impure para-xylene of preferred purity are alreadyknown in the art and include, but are not limited to, distillation,partial crystallization methods at sub-ambient temperatures, andmolecular sieve methods using selective pore-size adsorption. However,the preferred ranges of purity specified herein are, at their high end,more demanding and expensive than characteristically practiced bycommercial suppliers of para-xylene; and yet at the low end, thepreferred ranges avoid overly costly purification of para-xylene forfeeding to a partial oxidation reaction medium by discovering anddisclosing where the combined effects of impurity self-generation frompara-xylene itself and of impurity consumptive reactions within thereaction medium become more important than the feed rates of impuritieswithin impure para-xylene.

When the xylene-containing feed stream contains selected impurities,such as ethyl-benzene and/or toluene, oxidation of these impurities cangenerate benzoic acid. As used herein, the term “impurity-generatedbenzoic acid” shall denote benzoic acid derived from any source otherthan xylene during xylene oxidation.

As disclosed herein, a portion of the benzoic acid produced duringxylene oxidation is derived from the xylene itself. This production ofbenzoic acid from xylene is distinctly in addition to any portion ofbenzoic acid production that may be impurity-generated benzoic acid.Without being bound by theory, it is believed that benzoic acid isderived from xylene within the reaction medium when various intermediateoxidation products of xylene spontaneously decarbonylate (carbonmonoxide loss) or decarboxylate (carbon dioxide loss) to thereby producearyl radicals. These aryl radicals can then abstract a hydrogen atomfrom one of many available sources in the reaction medium and produceself-generated benzoic acid. Whatever the chemical mechanism, the term“self-generated benzoic acid”, as used herein, shall denote benzoic acidderived from xylene during xylene oxidation.

As also disclosed herein, when para-xylene is oxidized to produceterephthalic acid (TPA), the production of self-generated benzoic acidcauses para-xylene yield loss and oxidant yield loss. In addition, thepresence of self-generated benzoic acid in the liquid phase of thereaction medium correlates with increases for many undesirable sidereactions, notably including generation of highly colored compoundscalled mono-carboxy-fluorenones. Self-generated benzoic acid alsocontributes to the undesirable accumulation of benzoic acid in recycledfiltrate which further elevates the concentration of benzoic acid in theliquid phase of the reaction medium. Thus, formation of self-generatedbenzoic acid is desirably minimized, but this is also appropriatelyconsidered simultaneously with impurity-generated benzoic acid, withfactors affecting consumption of benzoic acid, with factors pertainingto other issues of reaction selectivity, and with overall economics.

The inventors have discovered that the self-generation of benzoic acidcan be controlled to low levels by appropriate selection of, forexample, temperature, xylene distribution, and oxygen availabilitywithin the reaction medium during oxidation. Not wishing to be bound bytheory, lower temperatures and improved oxygen availability appear tosuppress the decarbonylation and/or decarboxylation rates, thus avoidingthe yield loss aspect of self-generated benzoic acid. Sufficient oxygenavailability appears to direct aryl radicals toward other more benignproducts, in particular hydroxybenzoic acids. Distribution of xylene inthe reaction medium may also affect the balance between aryl radicalconversion to benzoic acid or to hydroxybenzoic acids. Whatever thechemical mechanisms, the inventors have discovered reaction conditionsthat, although mild enough to reduce benzoic acid production, are severeenough to oxidize a high fraction of the hydroxybenzoic acid productionto carbon monoxide and/or carbon dioxide, which are easily removed fromthe oxidation product.

In a preferred embodiment of the present invention, the oxidationreactor is configured and operated in a manner such that the formationof self-generated benzoic acid is minimized and the oxidation ofhydroxybenzoic acids to carbon monoxide and/or carbon dioxide ismaximized. When the oxidation reactor is employed to oxidize para-xyleneto terephthalic acid, it is preferred that para-xylene makes up at leastabout 50 weight percent of the total xylene in the feed streamintroduced into the reactor. More preferably, para-xylene makes up atleast about 75 weight percent of the total xylene in the feed stream.Still, more preferably, para-xylene makes up at least 95 weight percentof the total xylene in the feed stream. Most preferably, para-xylenemakes up substantially all of the total xylene in the feed stream.

When the reactor is employed to oxidize para-xylene to terephthalicacid, it is preferred for the rate of production of terephthalic acid tobe maximized, while the rate of production of self-generated benzoicacid is minimized. Preferably, the ratio of the rate of production (byweight) of terephthalic acid to the rate of production (by weight) ofself-generated benzoic acid is at least about 500:1, more preferably atleast about 1,000:1, and most preferably at least 1,500:1. As will beseen below, the rate of production of self-generated benzoic acid ispreferably measured when the concentration of benzoic acid in the liquidphase of the reaction medium is below 2,000 ppmw, more preferably below1,000 ppmw, and most preferably below 500 ppmw, because these lowconcentrations suppress to suitably low rates reactions that convertbenzoic acid to other compounds.

Combining the self-generated benzoic acid and the impurity-generatedbenzoic acid, the ratio of the rate of production (by weight) ofterephthalic acid to the rate of production (by weight) of total benzoicacid is preferably at least about 400:1, more preferably at least about700:1, and most preferably at least 1,100:1. As will be seen below, thesummed rate of production of self-generated benzoic acid plusimpurity-generated benzoic acid is preferably measured when theconcentration of benzoic acid in the liquid phase of the reaction mediumis below 2,000 ppmw, more preferably below 1,000 ppmw, and mostpreferably below 500 ppmw, because these low concentrations suppress tosuitably low rates reactions that convert benzoic acid to othercompounds.

As disclosed herein, elevated concentrations of benzoic acid in theliquid phase of the reaction medium lead to increased formation of manyother aromatic compounds, several of which are noxious impurities inTPA; and, as disclosed herein, elevated concentrations of benzoic acidin the liquid phase of the reaction medium lead to increased formationof carbon oxide gases, the formation of which represents yield loss onoxidant and on aromatic compounds and/or solvent. Furthermore, it is nowdisclosed that the inventors have discovered a considerable portion ofthis increased formation of other aromatic compounds and of carbonoxides derives from reactions that convert some of the benzoic acidmolecules themselves, as contrasted to benzoic acid catalyzing otherreactions without itself being consumed. Accordingly, the “netgeneration of benzoic acid” is defined herein as the time-averagedweight of all benzoic acid exiting the reaction medium minus thetime-averaged weight of all benzoic acid entering the reaction mediumduring the same period of time. This net generation of benzoic acid isoften positive, driven by the formation rates of impurity-generatedbenzoic acid and of self-generated benzoic acid. However, the inventorshave discovered that the conversion rate of benzoic acid to carbonoxides, and to several other compounds, appears to increaseapproximately linearly as the concentration of benzoic acid is increasedin the liquid phase of the reaction medium, measured when other reactionconditions comprising temperature, oxygen availability, STR, andreaction activity are maintained appropriately constant. Thus, when theconcentration of benzoic acid in the liquid-phase of the reaction mediumis great enough, perhaps due to an elevated concentration of benzoicacid in recycled solvent, then the conversion of benzoic acid moleculesto other compounds, including carbon oxides, can become equal to orgreater than the chemical generation of new benzoic acid molecules. Inthis case, the net generation of benzoic acid can become balanced nearzero or even negative. The inventors have discovered that when the netgeneration of benzoic acid is positive, then the ratio of the rate ofproduction (by weight) of terephthalic acid in the reaction mediumcompared to the rate of net generation of benzoic acid in the reactionmedium is preferably above about 700:1, more preferably above about1,100:1, and most preferably above 4,000:1. The inventors havediscovered that when the net generation of benzoic acid is negative, theratio of the rate of production (by weight) of terephthalic acid in thereaction medium compared to the rate of net generation of benzoic acidin the reaction medium is preferably above about 200:(−1), morepreferably above about 1,000:(−1), and most preferably above 5,000:(−1).

The inventors have also discovered preferred ranges for the compositionof the slurry (liquid+solid) withdrawn from the reaction medium and forthe solid CTA portion of the slurry. The preferred slurry and thepreferred CTA compositions are surprisingly superior and useful. Forexample, purified TPA produced from this preferred CTA by oxidativedigestion has a sufficiently low level of total impurities and ofcolored impurities such that the purified TPA is suitable, withouthydrogenation of additional 4-CBA and/or colored impurities, for a widerange of applications in PET fibers and PET packaging applications. Forexample, the preferred slurry composition provides a liquid phase of thereaction medium that is relatively low in concentration of importantimpurities and this importantly reduces the creation of other even moreundesirable impurities as disclosed herein. In addition, the preferredslurry composition importantly aids the subsequent processing of liquidfrom the slurry to become suitably pure recycled solvent, according toother embodiments of the present invention.

CTA produced according to one embodiment of the present inventioncontains less impurities of selected types than CTA produce byconventional processes and apparatuses, notably those employing recycledsolvent. Impurities that may be present in CTA include the following:4-carboxybenzaldehyde (4-CBA), 4,4′-dicarboxystilbene (4,4′-DCS),2,6-dicarboxyanthraquinone (2,6-DCA), 2,6-dicarboxyfluorenone (2,6-DCF),2,7-dicarboxyfluorenone (2,7-DCF), 3,5-dicarboxyfluorenone (3,5-DCF),9-fluorenone-2-carboxylic acid (9F-2CA), 9-fluorenone-4-carboxylic acid(9F-4CA), 4,4′-dicarboxybiphenyl (4,4′-DCB), 2,5,4′-tricarboxybiphenyl(2,5,4′-TCB), phthalic acid (PA), isophthalic acid (IPA), benzoic acid(BA), trimellitic acid (TMA), para-toluic acid (PTAC),2,6-dicarboxybenzocoumarin (2,6-DCBC), 4,4′-dicarboxybenzil (4,4′-DCBZ),4,4′-dicarboxybenzophenone (4,4′-DCBP), 2,5,4′-tricarboxybenzophenone(2,5,4′-TCBP). Table 3, below provides the preferred amounts of theseimpurities in CTA produced according to an embodiment of the presentinvention. TABLE 3 CTA Impurities Impurity Preferred More Preferred MostPreferred Identification Amt. (ppmw) Amt. (ppmw) Amt. (ppmw) 4-CBA<15,000   100-8,000   400-2,000 4,4′-DCS <12 <6 <3 2,6-DCA <9 <6 <22,6-DCF <100  2-50  5-25 2,7-DCF <30 <15 <5 3,5-DCF <16 <8 <2 9F-2CA <16<8 <4 9F-4CA <8 <4 <2 Total fluorenones <100  2-60  4-35 4,4′-DCB <64 1-32 2-8 2,5,4′-TCB <24 <12 <8 PA <200  3-100  5-50 IPA <800  10-400 20-200 BA <600  5-300  15-100 TMA <800  10-400  20-200 PTAC <2,000  10-1,000  50-500 2,6-DCBC <64 <32 <8 4,4′-DCBZ <12 <8 <4 4,4′-DCBP <40<30 <20 2,5,4′-TCBP <32 <16 <4

In addition, it is preferred for CTA produced according to an embodimentof the present invention to have reduced color content relative to CTAproduce by conventional processes and apparatuses, notably thoseemploying recycled solvent. Thus, it is preferred for CTA produced inaccordance to one embodiment of the present invention have a percenttransmittance percent at 340 nanometers (nm) of at least about 25percent, more preferably of at least about 50 percent, and mostpreferably of at least 60 percent. It is further preferred for CTAproduced in accordance to one embodiment of the present invention tohave a percent transmittance percent at 400 nanometers (nm) of at leastabout 88 percent, more preferably of at least about 90 percent, and mostpreferably of at least 92 percent.

The test for percent transmittance provides a measure of the colored,light-absorbing impurities present within TPA or CTA. As used herein,the test refers to measurements done on a portion of a solution preparedby dissolving 2.00 grams of dry solid TPA or CTA in 20.0 milliliters ofdimethyl sulfoxide (DMSO), analytical grade or better. A portion of thissolution is then placed in a Hellma semi-micro flow cell, PN 176.700,which is made of quartz and has a light path of 1.0 cm and a volume of0.39 milliliters. (Hellma USA, 80 Skyline Drive, Plainview, N.Y. 11803).An Agilent 8453 Diode Array Spectrophotometer is used to measure thetransmittance of different wavelengths of light through this filled flowcell. (Agilent Technologies, 395 Page Mill Road, Palo Alto, Calif.94303). After appropriate correction for absorbance from the background,including but not limited to the cell and the solvent used, the percenttransmittance results, characterizing the fraction of incident lightthat is transmitted through the solution, are reported directly by themachine. Percent transmittance values at light wavelengths of 340nanometers and 400 nanometers are particularly useful for discriminatingpure TPA from many of the impurities typically found therein.

The preferred ranges of various aromatic impurities in the slurry(solid+liquid) phase of the reaction medium are provided below in Table4. TABLE 4 Slurry Impurities Impurity Preferred More Preferred MostPreferred Identification Amt. (ppmw) Amt. (ppmw) Amt. (ppmw) 4-CBA<8,000 <5,000 <2,500 4,4′-DCS <4 <2 <1 2,6-DCA <6 <3 <1 2,6-DCF <70 2-40  4-20 2,7-DCF <12 <8 <4 3,5-DCF <12 <8 <4 9F-2CA <12 <8 <4 9F-4CA<8 <4 <2 Total fluorenones <90  2-60  5-30 4,4′-DCB <64  1-16 2-42,5,4′-TCB <60  2-40  4-20 PA <3,000   25-1,500  75-500 IPA 9,000  75-4,500   225-1,500 BA <15,000   100-6,000   300-2,000 TMA <3,000  25-1,500  75-500 PTAC <8,000   100-4,000   200-2,000 4,4′-DCBZ <5 <4<3 4,4′-DCBP <240 <160 <80 2,5,4′-TCBP <120 <80 <40

These preferred compositions for the slurry embody the preferredcomposition of the liquid phase of the reaction medium while usefullyavoiding experimental difficulties pertaining to precipitation ofadditional liquid phase components from the reaction medium into solidphase components during sampling from the reaction medium, separation ofliquids and solids, and shifting to analytical conditions.

Many other aromatic impurities are also typically present in the slurryphase of the reaction medium and in CTA of the reaction medium,generally varying at even lower levels and/or in proportion to one ormore of the disclosed aromatic compounds. Controlling the disclosedaromatic compounds in the preferred ranges will keep other aromaticimpurities at suitable levels. These advantaged compositions for theslurry phase in the reaction medium and for the solid CTA taken directlyfrom the slurry are enabled by operating with embodiments of theinvention disclosed herein for partial oxidation of para-xylene to TPA.

Measurement of the concentration of low level components in the solvent,recycled solvent, CTA, slurry from the reaction medium, and PTA areperformed using liquid chromatography methods. Two interchangeableembodiments are now described.

The method referred to herein as HPLC-DAD comprises high pressure liquidchromatography (HPLC) coupled with a diode array detector (DAD) toprovide separation and quantitation of various molecular species withina given sample. The instrument used in this measurement is a model 1100HPLC equipped with a DAD, provided by Agilent Technologies (Palo Alto,Calif.), though other suitable instruments are also commerciallyavailable and from other suppliers As is known in the art, both theelution time and the detector response are calibrated using knowncompounds present in known amounts, compounds and amounts that areappropriate to those occurring in actual unknown samples.

The method referred to herein as HPLC-MS comprises high pressure liquidchromatography (HPLC) coupled with mass spectrometry (MS) to provideseparation, identification, and quantitation of various molecularspecies within a given sample. The instruments used in this measurementis an Alliance HPLC and ZQ MS provided by Waters Corp. (Milford, Mass.),though other suitable instruments are also commercially available andfrom other suppliers. As is known in the art, both the elution time andthe mass spectrometric response are calibrated using known compoundspresent in known amounts, compounds and amounts that are appropriate tothose occurring in actual unknown samples.

Another embodiment of the current invention relates to partial oxidationof aromatic oxidizable compound with appropriate balancing of thesuppression of noxious aromatic impurities on the one hand against theproduction of carbon dioxide and carbon monoxide, collectively carbonoxides (COx), on the other. These carbon oxides typically exit thereaction vessel in the off-gas, and they correspond to a destructiveloss of solvent and of oxidizable compound, including the ultimatelypreferred oxidized derivatives (e.g., acetic acid, para-xylene, andTPA). The inventors have discovered lower bounds for the production ofcarbon oxides below which it seems the high creation of noxious aromaticimpurities, as described below, and the low overall conversion level areinevitably too poor to be of economic utility. The inventors have alsodiscovered upper bounds of carbon oxides above which the generation ofcarbon oxides continues to increase with little further value providedby reduction in generation of noxious aromatic impurities.

The inventors have discovered that reducing the liquid-phaseconcentrations of aromatic oxidizable compound feed and of aromaticintermediate species within a reaction medium leads to lower generationrates for noxious impurities during the partial oxidation of aromaticoxidizable compound. These noxious impurities include coupled aromaticrings and/or aromatic molecules containing more than the desired numberof carboxylic acid groups (e.g., in the oxidation of para-xylene thenoxious impurities include 2,6-dicarboxyanthraquinone,2,6-dicarboxyfluorenone, trimellitic acid, 2,5,4′-tricarboxybiphenyl,and 2,5,4′-benzophenone). The aromatic intermediate species includearomatic compounds descended from the feed of oxidizable aromaticcompound and still retaining non-aromatic hydrocarbyl groups (e.g., inthe oxidation of para-xylene the aromatic intermediate species comprisepara-tolualdehyde, terephthaldehyde, para-toluic acid, 4-CBA,4-hydroxymethylbenzoic acid, and alpha-bromo-para-toluic acid). Thearomatic oxidizable compound feed and the aromatic intermediate speciesretaining non-aromatic hydrocarbyl groups, when present in the liquidphase of the reaction medium, appear to lead to noxious impurities in amanner similar to that already disclosed herein for dissolved aromaticspecies lacking non-aromatic hydrocarbyl groups (e.g., isophthalicacid).

Set against this need for higher reaction activity to suppress formationof noxious aromatic impurities during partial oxidation of oxidizablearomatic compound, the inventors have discovered that the undesirableattendant result is increased production of carbon oxides. It isimportant to appreciate that these carbon oxides represent a yield lossof oxidizable compound and oxidant, not just solvent. Explicitly, asubstantial and sometimes principal fraction of the carbon oxides comesfrom the oxidizable compound, and its derivatives, rather than fromsolvent; and often the oxidizable compound costs more per carbon unitthan does solvent. Furthermore, it is important to appreciate that thedesired product carboxylic acid (e.g., TPA) is also subject toover-oxidation to carbon oxides when present in the liquid phase of thereaction medium.

It is also important to appreciate that the present invention relates toreactions in the liquid phase of the reaction medium and to reactantconcentrations therein. This is in contrast to some prior inventionswhich relate directly to the creation in precipitated solid form ofaromatic compound retaining non-aromatic hydrocarbyl groups.Specifically, for the partial oxidation of para-xylene to TPA, certainprior inventions pertain to the amount of 4-CBA precipitated in thesolid phase of CTA. However, the present inventors have discovered avariance of greater than two to one for the ratio of 4-CBA in the solidphase to 4-CBA in the liquid phase, using the same specifications oftemperature, pressure, catalysis, solvent composition and space-timereaction rate of para-xylene, depending upon whether the partialoxidation is conducted in a well-mixed autoclave or in a reaction mediumwith oxygen and para-xylene staging according to the present invention.Further, the inventors have observed that the ratio of 4-CBA in thesolid phase to 4-CBA in the liquid phase can also vary by over two toone in either well-mixed or staged reaction medium depending upon thespace-time reaction rate of para-xylene at otherwise similarspecifications of temperature, pressure, catalysis, and solventcomposition. Additionally, 4-CBA in the solid phase CTA does not appearto contribute to the formation of noxious impurities, and 4-CBA in thesolid phase can be recovered and oxidized on to TPA simply and at highyield (e.g., by oxidative digestion of the CTA slurry as is describedherein); whereas the removal of noxious impurities is far more difficultand costly than removal of solid phase 4-CBA, and the production ofcarbon oxides represents a permanent yield loss. Thus, it is importantto distinguish that this aspect of the present invention relates toliquid-phase compositions in the reaction medium.

Whether sourced from solvent or oxidizable compound, the inventors havediscovered that at conversions of commercial utility the production ofcarbon oxides relates strongly to the level of overall reaction activitydespite wide variation in the specific combination of temperature,metals, halogens, temperature, acidity of the reaction medium asmeasured by pH, water concentration employed to obtain the level ofoverall reaction activity. The inventors have found it useful for thepartial oxidation of xylene to evaluate the level of overall reactionactivity using the liquid-phase concentration of toluic acids at themid-height of the reaction medium, the bottom of the reaction medium,and the top of the reaction medium.

Thus, there arises an important simultaneous balancing to minimize thecreation of noxious impurities by increasing reaction activity and yetto minimize the creation of carbon oxides by lowering reaction activity.That is, if the overall production of carbon oxides is suppressed toolow, then excessive levels of noxious impurities are formed, and viceversa.

Furthermore, the inventors have discovered that the solubility and therelative reactivity of the desired carboxylic acid (e.g., TPA) and thepresence of other dissolved aromatic species lacking non-aromatichydrocarbyl groups introduce a very important fulcrum in this balancingof carbon oxides versus noxious impurities. The desired productcarboxylic acid is typically dissolved in the liquid phase of thereaction medium, even when also present in solid form. For example, attemperatures in the preferred ranges, TPA is soluble in a reactionmedium comprising acetic acid and water at levels ranging from about onethousand ppmw to in excess of 1 weight percent, with solubilityincreasing as temperature increases. Notwithstanding that there aredifferences in the reaction rates toward forming various noxiousimpurities from oxidizable aromatic compound feed (e.g., para-xylene),from aromatic reaction intermediates (e.g., para-toluic acid), from thedesired product aromatic carboxylic acid (e.g., TPA), and from aromaticspecies lacking non-aromatic hydrocarbyl groups (e.g., isophthalicacid), the presence and reactivity of the latter two groups establishesa region of diminishing returns with regards to further suppression ofthe former two groups, oxidizable aromatic compound feed and aromaticreaction intermediates. For example, in a partial oxidation ofpara-xylene to TPA, if dissolved TPA amounts to 7,000 ppmw in the liquidphase of the reaction medium at given conditions, dissolved benzoic acidamounts to 8,000 ppmw, dissolved isophthalic acid amounts to 6,000 ppmw,and dissolved phthalic acid amounts to 2,000 ppmw, then the value towardfurther lowering of total noxious compounds begins to diminish asreaction activity is increased to suppress the liquid-phaseconcentration para-toluic acid and 4-CBA below similar levels. That is,the presence and concentration in the liquid phase of the reactionmedium of aromatic species lacking non-aromatic hydrocarbyl groups isvery little altered by increasing reaction activity, and their presenceserves to expand upwards the region of diminishing returns for reducingthe concentration of reaction intermediates in order to suppressformation of noxious impurities.

Thus, one embodiment of the present invention provides preferred rangesof carbon oxides, bounded on the lower end by low reaction activity andexcessive formation of noxious impurities and on upper end by excessivecarbon losses, but at levels lower than previously discovered anddisclosed as commercially useful. Accordingly, the formation of carbonoxides is preferably controlled as follows. The ratio of moles of totalcarbon oxides produced to moles of oxidizable aromatic compound fed ispreferably greater than about 0.02:1, more preferably greater than about0.04:1, still more preferably greater than about 0.05:1, and mostpreferably greater than 0.06:1. At the same time, the ratio of moles oftotal carbon oxides produced to moles of oxidizable aromatic compoundfed is preferably less than about 0.24:1, more preferably less thanabout 0.22:1, still more preferably less than about 0.19:1, and mostpreferably less than 0.15:1. The ratio of moles of carbon dioxideproduced to moles of oxidizable aromatic compound fed is preferablygreater than about 0.01:1, more preferably greater than about 0.03:1,still more preferably greater than about 0.04:1, and most preferablygreater than 0.05:1. At the same time, the ratio of moles of carbondioxide produced to moles of oxidizable aromatic compound fed ispreferably less than about 0.21:1, more preferably less than about0.19:1, still more preferably less than about 0.16:1, and mostpreferably less than 0.11. The ratio of moles of carbon monoxideproduced to moles of oxidizable aromatic compound fed is preferablygreater than about 0.005:1, more preferably greater than about 0.010:1,still more preferably greater than about 0.015:1, and most preferablygreater than 0.020:1. At the same time, the ratio of moles of carbonmonoxide produced to moles of oxidizable aromatic compound fed ispreferably less than about 0.09:1, more preferably less than about0.07:1, still more preferably less than about 0.05:1, and mostpreferably less than 0.04:1

The content of carbon dioxide in dry off-gas from the oxidation reactoris preferably greater than about 0.10 mole percent, more preferablygreater than about 0.20 mole percent, still more preferably greater thanabout 0.25 mole percent, and most preferably greater than 0.30 molepercent. At the same time, the content of carbon dioxide in dry off-gasfrom the oxidation reactor is preferably less than about 1.5 molepercent, more preferably less than about 1.2 mole percent, still morepreferably less than about 0.9 mole percent, and most preferably lessthan 0.8 mole percent. The content of carbon monoxide in dry off-gasfrom the oxidation reactor is preferably greater than about 0.05 molepercent, more preferably greater than about 0.10 mole percent, stillmore preferably greater than 0.15, and most preferably greater than 0.18mole percent. At the same time, the content of carbon monoxide in dryoff-gas from the oxidation reactor is preferably less than about 0.60mole percent, more preferably less than about 0.50 mole percent, stillmore preferably less than about 0.35 mole percent, and most preferablyless than 0.28 mole percent

The inventors have discovered that an important factor for reducing theproduction of carbon oxides to these preferred ranges is improving thepurity of the recycled filtrate and of the feed of oxidizable compoundto reduce the concentration of aromatic compounds lacking non-aromatichydrocarbyl groups according to disclosures of the presentinvention—this simultaneously reduces the formation of carbon oxides andof noxious impurities. Another factor is improving distribution ofpara-xylene and oxidant within the reaction vessel according todisclosures of the present invention. Other factors enabling the abovepreferred levels of carbon oxides are to operate with the gradients inthe reaction medium as disclosed herein for pressure, for temperature,for concentration of oxidizable compound in the liquid phase, and foroxidant in the gas phase. Other factors enabling the above preferredlevels of carbon oxides are to operate within the disclosures hereinpreferred for space-time reaction rate, pressure, temperature, solventcomposition, catalyst composition, and mechanical geometry of thereaction vessel.

An important benefit from operating within the preferred ranges ofcarbon oxide formation is that the usage of molecular oxygen can bereduced, though not to stoichiometric values. Notwithstanding the goodstaging of oxidant and oxidizable compound according to the presentinvention, an excess of oxygen must be retained above the stoichiometricvalue, as calculated for feed of oxidizable compound alone, to allow forsome losses to carbon oxides and to provide excess molecular oxygen tocontrol the formation of noxious impurities. Specifically for the casewhere xylene is the feed of oxidizable compound, the feed ratio ofweight of molecular oxygen to weight of xylene is preferably greaterthan about 0.91:1.00, more preferably greater than about 0.95:1.00, andmost preferably greater than 0.99:1.00. At the same time, the feed ratioof weight of molecular oxygen to weight of xylene is preferably lessthan about 1.20:1.00, more preferably less than about 1.12:1.00, andmost preferably less than 1.06:1.00. Specifically for xylene feed, thetime-averaged content of molecular oxygen in the dry off-gas from theoxidation reactor is preferably greater than about 0.1 mole percent,more preferably greater than about 1 mole percent, and most preferablygreater than 1.5 mole percent. At the same time, the time-averagedcontent of molecular oxygen in the dry off-gas from the oxidationreactor is preferably less than about 6 mole percent, more preferablyless than about 4 mole percent, and most preferably less than 3 molepercent.

Another important benefit from operating within the preferred ranges ofcarbon oxide formation is that less aromatic compound is converted tocarbon oxides and other less valuable forms. This benefit is evaluatedusing the sum of the moles of all aromatic compounds exiting thereaction medium divided by the sum of the moles of all aromaticcompounds entering the reaction medium over a continuous period of time,preferably one hour, more preferably one day, and most preferably 30consecutive days. This ratio is hereinafter referred to as the “molarsurvival ratio” for aromatic compounds through the reaction medium andis expressed as a numerical percentage. If all entering aromaticcompounds exit the reaction medium as aromatic compounds, albeit mostlyin oxidized forms of the entering aromatic compounds, then the molarsurvival ratio has its maximum value of 100 percent. If exactly 1 ofevery 100 entering aromatic molecules is converted to carbon oxidesand/or other non-aromatic molecules (e.g., acetic acid) while passingthrough reaction medium, then the molar survival ratio is 99 percent.Specifically for the case where xylene is the principal feed ofoxidizable aromatic compound, the molar survival ratio for aromaticcompounds through the reaction medium is preferably greater than about98 percent, more preferably greater than about 98.5 percent, and mostpreferably less than 99.0 percent. At the same time and in order thatsufficient overall reaction activity is present, the molar survivalratio for aromatic compounds through the reaction medium is preferablyless than about 99.9 percent, more preferably less than about 99.8percent, and most preferably less than 99.7 percent when xylene is theprincipal feed of oxidizable aromatic compound.

Another aspect of the current invention involves the production ofmethyl acetate in a reaction medium comprising acetic acid and one ormore oxidizable aromatic compounds. This methyl acetate is relativelyvolatile compared to water and acetic acid and thus tends to follow theoff-gas unless additional cooling or other unit operations are employedto recover it and/or to destroy it prior to releasing the off-gas backto the environment. The formation of methyl acetate thus represents anoperating cost and also a capital cost. Perhaps the methyl acetate isformed by first combining a methyl radical, perhaps from decompositionof acetic acid, with oxygen to produce methyl hydroperoxide, bysubsequently decomposing to form methanol, and by finally reacting theproduced methanol with remaining acetic acid to form methyl acetate.Whatever the chemical path, the inventors have discovered that whenevermethyl acetate production is at too low a rate, then the production ofcarbon oxides are also too low and the production of noxious aromaticimpurities are too high. If methyl acetate production is at too high arate, then the production of carbon oxides are also unnecessarily highleading to yield losses of solvent, oxidizable compound and oxidant.When employing the preferred embodiments disclosed herein, theproduction ratio of moles of methyl acetate produced to moles ofoxidizable aromatic compound fed is preferably greater than about0.005:1, more preferably greater than about 0.010:1, and most preferablygreater than 0.020:1. At the same time, the production ratio of moles ofmethyl acetate produced to moles of oxidizable aromatic compound fed ispreferably less than about 0.09:1 more preferably less than about0.07:1, still more preferably less than about 0.05:1, and mostpreferably less than 0.04:1.

This invention can be further illustrated by the following examples ofpreferred embodiments thereof, although it will be understood that theseexamples are included merely for purposes of illustration and are notintended to limit the scope of the invention unless otherwisespecifically indicated.

EXAMPLE 1

This is an operational example from a commercial oxidation ofpara-xylene in a bubble column reactor. This example demonstrates, forexample, that large vertical gradients exist for concentrations ofpara-xylene when appropriate geometric and process conditions areemployed according to aspects of the current invention.

This example employed a commercial bubble column oxidizer vessel havinga nearly vertical, essentially cylindrical body with an inside diameterof about 2.44 meters. The height of the bubble column oxidizer vesselwas about 32 meters from lower tangent line (TL) to upper TL. The vesselwas fitted with about 2:1 elliptical heads at the top and bottom of thecylinder. The operating level was about 25 meters of reaction mediumabove the lower TL. The feed rate of commercial-purity para-xylene waseffectively steady at a rate of about 81 kilograms per minute, enteringthe reaction vessel through a circular hole located in the wall of thecylindrical section at an elevation of about 4.35 meters above the lowerTL. The internal diameter of said wall hole was about 0.076 meters. Afiltrate solvent was fed at an effectively steady rate of about 777kilograms per minute. An unmetered fraction of this filtrate solvent,estimated from conduit sizes and pressure drops to be about 20 kilogramsper minute, was feed as a liquid flush to the oxidant sparger. Thebalance of the filtrate solvent, about 757 kilograms per minute, was fedintimately commingled with the commercial-purity para-xylene. Thecombined liquid-phase feed stream of filtrate solvent andcommercial-purity para-xylene thus amounted to about 838 kilograms perminute giving a superficial velocity of the inlet flow through said wallhole of about 3 meters per second. This filtrate solvent was from aplant recycle system and was comprised above about 97 weight percent ofacetic acid and water. The concentration of catalyst components in thefiltrate solvent was such that the composition within the liquid phaseof the reaction medium was about 1,777 ppmw of cobalt, about 1,518 ppmwof bromine, and about 107 ppmw of manganese. A separate stream of refluxsolvent was fed as droplets into the gas-disengaging zone above theoperating level of the reaction medium at an effectively steady rate ofabout 572 kilograms per minute. This reflux solvent was comprised ofabove about 99 weight percent of acetic acid and water; and the refluxsolvent was from a separate plant recycle system that was withoutsignificant levels of catalyst components. The combined water content ofthe filtrate solvent feed and of the reflux solvent feed was such thatthe concentration of water within the liquid phase of the reactionmedium was about 6.0 weight percent. The oxidant was compressed air fedat an effectively steady rate of about 384 kilograms per minute throughan oxidant sparger similar to the one illustrated in FIGS. 2-5. Thisoxidant sparger comprised a mitered flow conduit that was approximatelyan equal-sided octagon with a crossing member connecting from one sideto the opposite side and traversing through the vertical axis ofsymmetry of the reaction vessel. The mitered flow conduit was made fromnominal 12-inch Schedule 10S piping components. The width of the octagonfrom the centroid of one side of the flow conduit to the centroid of theopposite side was about 1.83 meters. The octagon lay approximatelyhorizontal, and the mid-elevation of the octagonal conduit was about0.11 meters above the lower TL of the reaction vessel. The oxidantsparger contained 75 circular holes that were about 0.025 meters indiameter. The holes were situated approximately uniformly around theoctagon and cross member, lying near the top of said 12-inch piping.There was one circular hole with diameter of about 0.012 meters near thebottom of one side only of the octagonal conduit. The operating pressurein the reaction vessel overhead gas was steadily about 0.52 megapascalgauge. The reaction was operated in a substantially adiabatic manner sothat the heat of reaction elevated the temperature of the incoming feedsand evaporated much of the incoming solvent. Measured near themid-elevation of the reaction medium, the operating temperature wasabout 160° C. An exiting slurry comprising crude terephthalic acid (CTA)was removed from near the bottom of the lower elliptical head of thereaction vessel at an effectively steady rate. The flow rate of theexiting slurry was about 408 kilograms per minute.

Samples of slurry from the reaction medium were obtained from threeelevations in the reaction vessel, as described below. In determiningthe concentration of various species at various locations within thereaction medium, it was necessary to account for the stochastic natureof the system by taking enough samples to determine a time-averagedvalue of sufficient resolution.

One set of five samples was obtained from the exiting slurry conduitfrom near the bottom of the lower elliptical head of the reactionvessel. Another set of five samples was obtained from a wall holelocated at an elevation of about 12.4 meters above the lower TL of thereaction vessel. The third set of five samples was obtained from a wallhole located at an elevation of about 17.2 meters above the lower TL ofthe reaction vessel.

All slurry samples were analyzed by a calibrated gas chromatography (GC)method for composition of para-xylene and para-tolualdehyde in theliquid phase. Table 5, below, shows the average of the five results thatwere obtained from the three different column elevations. Results arereported as mass parts of analyte per million mass parts (ppmw) ofliquid phase. TABLE 5 para-xylene para-tolualdehyde Sample Location(ppmw) (ppmw) Side wall about 17.2 meters 21 140 Side wall about 12.4meters 63 317 Underflow from bottom head 455 960

These results show large gradients occurred vertically in the localconcentrations of para-xylene and para-tolualdehyde. For example, thegradient in concentration of para-xylene observed in data of Table 5 wasover 20:1 (455:21). These results demonstrate that the inherent fluidmixing of the entering para-xylene feed within the bubble column wasimportantly slower than the inherent reaction rates. To a lesser extent,vertical gradients also were observed for the concentrations of otherrelated aromatic reactive species in the reaction medium (e.g.,para-toluic acid and 4-carboxy benzaldehyde).

As is demonstrated in subsequent examples, detailed calculational modelsshow that the actual range of para-xylene concentration within theliquid phase of the reaction medium of this example was well in excessof 100:1. Even without executing a rigorous calculational model, thoseskilled in the art will recognize that the actual maximum concentrationof para-xylene occurred in the region near where the feed para-xylenewas introduced to the bubble column reaction vessel through the vesselwall. This elevation of maximum para-xylene concentration is about 4.35meters above the lower TL, in between the samples take from about 12.4meters and from the underflow. Similarly, the actual minimumconcentration of para-xylene likely occurred at or very near the top ofthe reaction medium at about 25 meters, well above the highest elevationfrom where the above samples were taken.

Concentrations of para-xylene and other oxidizable compounds can bemeasured for other locations within the reaction medium by employingsuitable mechanical devices for sampling at any position vertically orhorizontally within the reaction medium. Optionally, concentrations forpositions not physically sampled and chemically analyzed may becalculated with reasonable accuracy using computational models ofsufficient intricacy to cope with the highly complex fluid flowpatterns, chemical reaction kinetics, energy balance, vapor-liquid-solidequilibriums, and inter-phase exchange rates.

EXAMPLES 2-5

Examples 2-5 are calculational models of bubble column reactors eitheridentical to the reactor of Example 1 or generally similar withspecified improvements. The computational fluid dynamics (CFD) modelingperformed to generate Examples 2-5 was performed in accordance with themodeling method disclosed in co-pending U.S. Pat. App. Ser. No.60/594,774 entitled “Modeling of Liquid-Phase Oxidation,” the entiredisclosure of which is expressly incorporated herein by reference.

In Examples 2-5, the CFD modeling is performed using CFX release 5.7(ANSYS, Inc. 275 Technology Drive, Canonsburg, Pa. 15317). Examples 2-5comprise upwards of about 100,000 discrete spatial computational cellseach. Time steps useful in Examples 2-5 are less than 0.1 seconds.Multiple bubble sizes ranging in diameter from about 0.005 to about 0.20meters prove useful to tune the CFD model to approximate closely to theaverage bubble hold-up assessed via differential pressure measurement,to the vertical bubble hold-up profile assessed via gamma-scanning, andto the horizontal profiles of bubble hold-up assessed via computedtomography (CT) scans. To select appropriate bubble sizes andpopulations in the CFD models of Example 2-5, actual plant operatingdata was obtained for slurry bubble columns with cylindrical insidediameters of about 2.44 meters and about 3.05 meters operating with thereaction medium near the pertinent composition and process conditions asdisclosed below. The reference data for overall bubble hold-up wereobtained using differential pressures measured from near the base of thevessel and up to the overhead off-gas. The reference data for verticalbubble-hold-up profile were obtained using a gamma-emitting radioactivesource and detection method incremented up the outside of the reactionvessel in steps ranging from about 0.05 meters to about 0.3 meters. Thereference data for horizontal bubble hold-up profiles were obtained byCT scans performed on a nine by nine grid across a horizontal plane ofthe operating bubble column using a gamma-emitting radioactive sourceand detection method. That is, the source was positioned at a givenelevation at nine different positions spaced about equally around theperimeter of the bubble column. For each position of the gamma-radiationsource, the amount of gamma-radiation passing through the reactionvessel and reaction medium was detected at nine different positionsspaced about equally around the perimeter of the bubble column. Variousmathematical models were then applied to this discrete data to produceestimations of the variation of bubble hold-up throughout the reactionmedium for said elevation. Multiple horizontal CT scans were obtained ontwo different days, for two different elevations, and with two differentfeed rates of para-xylene, compressed air, etc.

The chemical reaction model for consumption of para-xylene in thisenvironment is tuned to match the reactant profiles for para-xylene asfound in Example 1 along with other data for similar temperatures,pressures, reaction intensities, catalysis, water concentration, and soon, from both commercial and pilot scale testing. As an indicativeapproximation, the pseudo-first order time constant for decay ofpara-xylene reactive tracer is equal to about 0.2 reciprocal seconds forabout 160° C. and about the mean conditions of the reaction medium usedin Examples 24.

Importantly, the CFD models of flow fields obtained in Examples 2-4produce large scale fluctuations in bubble swarms and liquid surges thatare generally consistent with the observed low frequency undulation inthe operating bubble column reaction vessel.

EXAMPLE 2

This example develops calculations pertinent to the mechanicalconfiguration of Example 1 and sets a comparative basis for Examples 3and 4. In this example, the mechanical configuration of the bubblecolumn reactor is identical to Example 1, having a 0.076-meter circulardiameter entry hole through the reaction vessel wall for the feed streamcomprising para-xylene and filtrate solvent. The feed rate ofpara-xylene is about 1.84 kilograms per second, higher than inExample 1. The feed rate of filtrate solvent fed intimately commingledwith the para-xylene is about 18.4 kilograms per second. The superficialvelocity of the combined stream of para-xylene plus filtrate solvententering through the wall hole is thus about 4 meters per second. Thefeed rate of reflux solvent in to the gas disengaging head space is 12.8kilograms per second. The feed rate of compressed air through theoxidant sparger is about 9 kilograms per second. The solids content ofthe reaction slurry is about 31 weight percent. The product slurry iswithdrawn from the center of the bottom head of the reaction vesselusing an effectively steady rate to maintain an approximately steadylevel of about 25 meters of reaction medium. The average gas hold-up forthe mid-elevation of the reaction medium is about 55 percent on anarea-averaged, time-averaged basis, where the length of time-averagingis at least about 100 seconds of CFD model time. The pressure in theheadspace above the reaction medium is about 0.50 megapascal gauge. Thetemperature is about 160° C. measured near the mid-elevation of thereaction medium. The contents of water and of cobalt, bromine, andmanganese within the liquid portion of the reaction medium areessentially the same as in Example 1.

EXAMPLE 3

This example develops calculations pertinent to improving dispersion ofpara-xylene feed by increasing the superficial velocity of theliquid-phase feed comprising para-xylene at its point of entry to thereaction medium according to one aspect of the current invention. Inthis example, the mechanical configuration of the bubble column reactoris identical to Example 2 except that the wall hole through which theliquid-phase feed comprising para-xylene is admitted is reduced to a0.025 meter circular diameter. The feed rate of para-xylene and otherprocess conditions are the same as for Example 2, excepting that thesuperficial velocity of the combined liquid-phase feed stream ofpara-xylene plus filtrate solvent entering through the wall hole is nowabout 36 meters per second.

The CFD model calculations of time-averaged fractions of reaction mediumwith para-xylene reactive tracer concentration in liquid phase abovevarious thresholds are presented in Table 6, below. The volume ofreaction medium with very highly concentrated para-xylene reactivetracer in the liquid phase is decreased by operating with higher inletvelocities of the liquid-phase feed stream comprising para-xyleneaccording to the present invention. The reduced regions of highpara-xylene concentration are important to limit undesirable couplingreactions both because concentrations of many soluble aromatic speciesare therein elevated and because such concentrations lead to locallyhigh consumption of dissolved molecular oxygen and thereby lead tolocally suppressed standing concentrations of dissolved molecularoxygen. TABLE 6 Example 2 Example 3 Example 4 Wall hole diameter(meters) 0.076 0.025 distributor Superficial velocity of incoming 4 36varies >15 pX + filtrate (m/sec) Percentage of reaction medium with pXconcentration above 1,000 ppmw (%) 3.64 3.28 3.73 above 2,500 ppmw (%)0.640 0.378 0.130 above 10,000 ppmw (%) 0.049 0.022 0.005 above 25,000ppmw (%) 0.009 0.002 0.001 Volume of reaction medium with pXconcentration above 1,000 ppmw (liters) 4,250 3,840 4,360 above 2,500ppmw (liters) 749 444 152 above 10,000 ppmw (liters) 57 26 6 above25,000 ppmw (liters) 10 2 1

EXAMPLE 4

This example develops calculations for improved mechanical means forintroducing oxidant and para-xylene into the bubble column reactor. Thisexample is executed within the same bubble column reactor as used inExamples 1-3. However, the reactor is modified with respect to themanner in which both the oxidant and the para-xylene are admitted intothe reaction medium. In discussing Example 4, attention is firstdirected to the modified apparatus for admitting para-xylene to thereaction medium, thereby reducing zones of high concentrations ofpara-xylene. Secondly, attention is directed to the modified apparatusfor admitting the oxidant to the reaction medium, thereby reducing zonesthat are poorly aerated. This is not to suppose that the twomodifications are totally independent in their results, but it is simplya step-wise presentation.

The amount of reaction medium with very high liquid phase concentrationsof para-xylene reactive tracer is reduced in Example 4 by use of aliquid-phase feed distribution system generally as shown in FIGS. 9-11.This liquid-phase feed distribution system conveniently has four flowconduits conveniently standing approximately vertical. Each of thesefour flow conduits is about 0.75 meters from the vertical axis ofsymmetry of the bubble column. These four flow conduits are convenientlymade from nominal 1.5-inch Schedule 10S piping components. The lower endof each leg in this example conveniently has a conically convergingsection with an included angle measured between opposite sides of thecone that is conveniently about 24 degrees; however, other shapes arealso useful to close the downstream end of the flow conduit (e.g. aconical closure with different included angle, a flat plate closure, apipe cap closure, a wedge-shaped closure, and so on.) Each of these fourflow conduits has a total of nine holes with each having a circulardiameter of about 0.0063 meters. The lowest one of the nine holes ineach conduit is at the bottom of the lower conical section. For eachconduit, this lowest hole is located about 0.4 meters above the lower TLof the reaction vessel. Measuring always from this bottom end of thetruncated bottom conical section, the next three holes in each conduitare elevated about 0.3 meters, the next three holes are elevated about1.6 meters, and the topmost two holes are elevated about 2.7 meters.Thus, the vertical distance from lowest hole to highest hole in eachconduit is about 2.7 meters, or about 1.1D. The linear (not vertical)distance of farthest hole separation, from the bottom hole of onevertical conduit to the top hole of the vertical conduit diagonallyopposite, is about 3.44 meters, or about 1.4D. For each level, the holesare spaced about evenly around the circumference of each flow conduit.The supply conduit for the feed of oxidizable compound and solvent tothe top of the four approximately vertical conduits is convenientlyabout horizontal at an elevation about 3.60 meters above the lower TL ofthe reaction vessel. The supply conduit is conveniently made fromnominal 3-inch Schedule 10S piping components. There is appropriatemechanical cross-bracing within the assembly and mechanical bracing fromthe assembly to the oxidant sparger and to the reaction vessel in orderto endure both static and dynamic forces occurring during both normaland upset operations.

Although not calculated in this example, many other designs for thisliquid-phase feed distribution system are possible. For example, theliquid flow conduit sizes can be larger or smaller or of differentcross-section than approximately circular or of different count thanfour. For example, each of the four essentially vertical conduits couldbe fed independently via flow conduits separately traversing thepressure containing wall of the reaction vessel. For example, theconnection to the supply of incoming para-xylene and feed solvent couldcome in near the mid-elevation or near the bottom elevation or at anyelevation or at multiple elevations of the approximately verticalconduits. For example, the supply conduits could be approximatelyvertical with the distribution holes residing in approximatelyhorizontal conduits, or both flow directions could be skewed ornon-linear or non-orthogonal. For example, the holes could be locateddifferently radially, azimuthally, or vertically with respect to thereaction medium. For example, more or fewer holes and/or holes ofdifferent shapes and/or holes with mixed sizes and/or mixed shapes canbe used. For example, exit nozzles could be used rather than exit holes.For example, one or more flow deflection apparatus can lie outside ofthe flow conduit close to the exit holes and in path of fluids uponexiting into the reaction medium.

Depending upon the solids character and content, if any, of the combinedfeed of para-xylene and solvent, or of the reaction medium, anddepending upon the start-up, shutdown and other operating proceduresemployed in actual manufacturing operation, it may be necessary to purgesolids from inside the liquid-phase feed distribution system. Althoughnot calculated in this example, a purging hole may usefully be largerthan the uniformly sized holes shown in the current example. The hole atthe lower end of each of the four approximately vertical legs isparticularly useful for purging solids, although it is not the onlypossible means. More complicated mechanical devices such as flapperassemblies, check valves, excess flow valves, power operated valves andthe like may be used either to prevent ingress of solids or to dischargeaccumulated solids from within the liquid-phase feed distributionsystem..

Now, attention is directed to the oxidant sparger, which is generally asshow in FIGS. 12-15. This oxidant sparger ring member convenientlycomprises a mitered flow conduit that is conveniently and approximatelyan equal-sided octagon without a crossing member. The mitered flowconduit is conveniently made from nominal 10-inch Schedule 10S pipingcomponents. The width of the octagon from the centroid of one side ofthe flow conduit to the centroid of the opposite side is about 1.12meters. The octagonal section conveniently lies approximatelyhorizontal, and the mid-elevation of the octagonal section is about 0.24meters below the lower TL of the reaction vessel. This is in distinctcontrast to the oxidant sparger ring member of Examples 1-3, theelevations of which are centered above the lower TL of the reactionvessel. The octagonal portion of the conduit is perforated with 64 aboutcircular holes each about 0.030 meters in diameter, approximatelyequally spaced around the conduit. About one-half of the holes arelocated around the conduit with locations that are at an angle of about45 degrees below horizontal, measuring from each hole to the nearestcentroid of the flow conduit cross-section. About one-half of the holesare located around the conduit with locations that are about at thebottom of the flow conduit (i.e., at an angle of about 90 degrees belowhorizontal, measuring from each hole to the nearest centroid of the flowconduit cross-section). The inventors again comment, akin to commentsfor the liquid-phase inlet distributor, that many other particulardesigns are possible for an oxidant sparger falling within the scope ofseveral aspects of the present invention. For example, more or less thantwo supply conduits may transverse the pressure containing wall. Forexample, the supply conduits of the oxidant sparger may be designedwithout comprising a ring member. For example, more than one ring membermay be present, and any ring member may have other than 8 sides or mayhave non-symmetrical sides. For example, the design may obtain apreferred pressure drop or a preferred quality of aeration or apreferred non-fouling nature while using a different number or size orsizes or placement of conduit holes or exits. For example, the designmay employ different diameters of conduits within preferred ranges. Forexample, the design of may achieve a non-fouling nature by using aliquid flush.

In this example, reaction medium is withdrawn with an effectively steadyrate from the side of the reaction vessel at an elevation of about 14meters through a wall hole that has an inside circular diameter of about0.076 meters. The withdrawn reaction medium is separated into a productslurry comprising crude terephthalic acid and an off-gas by using anexternal de-aeration vessel, which is described fully in Example 6. Theseparated off-gas from the external de-aeration vessel is conveyed by aconduit to join the main flow of off-gas leaving the top of the reactionvessel.

The CFD modeling methods of this example are substantially the same asfor Examples 2 and 3, with these exceptions. The spatial meshing ismodified as appropriate and known in the art for the revised apparatusfor distributing incoming oxidant, for distributing incoming oxidizablecompound, and for removing product slurry from the side wall of thereaction vessel about 14 meters above the lower TL.

To evaluate the results of the CFD model with respect to distribution ofthe para-xylene reactive tracer, the same methods are used as inExamples 2 and 3. Namely, the time-averaged fractions of reaction mediumwith para-xylene reactive tracer concentration in liquid phase abovevarious thresholds are determined. For ease in comparison, the resultsof this example are presented in Table 6, above. These results show thatimproved distribution of para-xylene reactive tracer of this exampleactually causes a small rise in the amount of reaction medium above1,000 ppmw, but the more harmful threshold levels of 2,500 ppmw, 10,000ppmw and 25,000 ppmw are reduced. These improvements are provided by,for example, higher feed inlet velocities along with improved vertical,radial and azimuthal positioning and spacing of the para-xyleneintroduction to the reaction medium.

Now turning to the quality of aeration throughout the reaction medium,the method of 2,000 horizontal slices of equal sub-volume is used toevaluate the amount of poorly aerated volume within the reaction mediumof Examples 24. Beginning at the lowest portion of the reaction medium,namely at the bottom of the lower elliptical head in this example, thereaction medium is partitioned into 2,000 equal sub-volumes usingtheoretical horizontal planes. For each of the CFD model timeincrements, within each of said 2,000 equal sub-volumes, the amount ofslurry and the amount of gas are determined and used to compute theaverage gas hold-up therein. To allow for the stochastic nature of theprocess, and of the CFD model thereof, the output from the CFD model istime-averaged through model times lasting at least about 100 seconds toobtain time-averaged values of gas-hold up in each of the 2,000 equalsub-volumes.

Once the time-averaged gas hold-up is determined for each of the 2,000equal sub-volumes, these values are compared to the threshold valuesdisclosed herein. For each threshold, the total number of offendingsub-volumes, those not exceeding the specified threshold value, areaccounted. Table 7, below, shows the number of 2,000 horizontal equalvolume slices of reaction medium with time-averaged gas hold-up below 10volume percent, below 20 volume percent, and below 30 volume percent forboth Example 2 and Example 4. Example 4 is importantly improved comparedto Example 2. TABLE 7 Example 2 Example 4 Elevation of centroid ofoxidant sparger ring +0.11 −0.24 member from lower TL of vessel (meters)Number of 2,000 horizontal equal volume slices of reaction medium withtime-averaged gas hold up below 10 volume percent 7 none below 20 volumepercent 21 none below 30 volume percent 41 none

In comparing calculational Examples 2 and 4, it is also notable that thepara-xylene feed of Example 4 is released lower in the reaction mediumand closer to the incoming oxidant stream than in Example 2.

EXAMPLES 5 and 6

Examples 5 and 6 are operational examples demonstrating in a commercialbubble column oxidizer the importance of minimizing regions of pooraeration, of improving the manner of introducing the commercial-puritypara-xylene feed to be more disperse vertically, azimuthally, andradially, and of lowering the entry of commercial-purity para-xylenefeed to be closer to the point of highest availability of molecularoxygen, according to the disclosures of the current invention.Additionally, these examples demonstrate a yield benefit from having anelevated slurry outlet.

There are many different impurity compounds typically produced by thecoupling of aromatic rings during the partial oxidation of para-xylene.One of these is 4,4′-dicarboxystilbene. This compound has a much higherabsorption of light than terephthalic acid has, and it strongly reducesthe optical transmittance of the intended product. In addition,4,4′-dicarboxystilbene is a convenient impurity to use in monitoring thequality of a continuous oxidation because it partitions selectively tothe solid phase of the reaction medium; therefore, very little4,4′-dicarboxystilbene is typically present in the recycle solventstreams of the commercial bubble column reaction vessels disclosed inExamples 5 and 6. In Examples 5 and 6, the concentrations of4,4′-dicarboxystilbene were measured with an analytical method employingHPLC-MS calibrated with a suitable reference mixture comprising solventand known amounts of several analytes specifically including a knownamount of 4,4′-dicarboxystilbene. The HPLC-MS analytical method isdescribed in above Detailed Description section.

EXAMPLE 5

The bubble column reactor employed in this example has substantially thesame mechanical configuration as the reactor of Examples 1 and 2. Thereactor is at process conditions comparable to Example 6 and provides acomparative basis. The operating level was about 25 meters of reactionmedium. The feed of commercial-purity para-xylene was effectively steadyat a rate of about 81 kilograms per minute. A filtrate solvent was fedat an effectively steady rate of about 793 kilograms per minute. Anunmetered fraction of this, estimated from conduit sizes and pressuredrops to be about 20 kilograms per minute, was feed as a liquid flush tothe oxidant sparger. The balance of the filtrate solvent, about 773kilograms per minute, was fed intimately commingled with thecommercial-purity para-xylene. The combined liquid-phase stream offiltrate solvent and commercial-purity para-xylene thus amounted toabout 854 kilograms per minute. This filtrate solvent was from a plantrecycle system and was comprised of above about 97 weight percent ofacetic acid and water. The concentration of catalyst components in thefiltrate solvent was such that the composition within the liquid phaseof the reaction medium was about 2,158 ppmw of cobalt, about 1,911 ppmwof bromine, and about 118 ppmw of manganese. A separate stream of refluxsolvent was fed as droplets into the gas-disengaging zone above theoperating level of the reaction medium at an effectively steady rate ofabout 546 kilograms per minute. This reflux solvent was comprised ofabove about 99 weight percent of acetic acid and water; and the refluxsolvent was from a separate plant recycle system that was withoutsignificant levels of catalyst components. The combined water content ofthe filtrate solvent feed and of the reflux solvent feed was such thatthe concentration of water within the liquid phase of the reactionmedium was about 5.8 weight percent. The oxidant was compressed air fedat an effectively steady rate of about 352 kilograms per minute. Theoperating pressure in the reaction vessel overhead gas was steadilyabout 0.42 megapascal gauge. The reaction vessel was operated in asubstantially adiabatic manner so that the heat of reaction elevated thetemperature of the incoming feeds and evaporated much of the incomingsolvent. Measured near the mid-elevation of the reaction medium, theoperating temperature was about 154.6° C. An exiting slurry comprisingcrude terephthalic acid (CTA) was removed from near the bottom of thelower elliptical head of the reaction vessel at an effectively steadyrate of about 428 kilograms per minute.

In this example, the ratio of the production rate of undesirable of4,4′-dicarboxystilbene to the production rate of desired terephthalicacid was measured by HPLC-MS on three separate samples of slurry productas about 8.6, 9.1, and 9.2 ppmw, thus averaging about 9.0 ppmw. Theconcentration of para-xylene in the liquid phase of the exiting slurrywas measured by calibrated GC on three separate samples of slurryproduct as about 777, 539, and 618 ppmw, thus averaging about 645 ppmw.The concentration of para-tolualdehyde in the liquid phase of theexiting slurry was measured by calibrated GC on said separate samples ofslurry product as about 1,055, 961, and 977 ppmw, thus averaging about998 ppmw.

EXAMPLE 6

The bubble column reactor of this example corresponds to the mechanicalconfiguration developed in calculational Example 4. The reactor of thisexample comprises improvements in the elevation, velocity, number andspacing of para-xylene feed entries, thus providing improveddistribution of para-xylene feed and improved staging against molecularoxygen. It further comprises improvements in the quality of aerationwithin the reaction medium, by using an improved oxidant sparger, and inthe elevation and method for removing and de-aerating slurry exiting thereaction medium. Compared to Example 5, important improvements are seenfor para-xylene yield, and important reductions are seen for impurityproduction.

The reactor of this example had the improved mechanical configuration asdescribed in CFD model Example 4. The operating level was about 25meters of reaction medium. The feed of commercial-purity para-xylene waseffectively steady at a rate of about 81 kilograms per minute. Afiltrate solvent was fed intimately commingled with thecommercial-purity para-xylene at an effectively steady rate of about 744kilograms per minute. The combined stream of filtrate solvent andcommercial-purity para-xylene feed thus amounted to about 825 kilogramsper minute. This filtrate solvent was from the same plant recycle systemand of substantially the same composition as in Example 5. Theconcentration of catalyst components in the filtrate solvent was suchthat the composition within the liquid phase of the reaction medium wasabout 1,996 ppmw of cobalt, about 1,693 ppmw of bromine, and about 108ppmw of manganese. A separate stream of reflux solvent was fed asdroplets into the gas-disengaging zone above the operating level of thereaction medium at an effectively steady rate of about 573 kilograms perminute. This reflux solvent was comprised of above about 99 weightpercent of acetic acid and water; and the reflux solvent was from aseparate plant recycle system that was without significant levels ofcatalyst components. The combined water content of the filtrate solventfeed and of the reflux solvent feed was such that the concentration ofwater within the liquid phase of the reaction medium was about 5.7weight percent. The oxidant was compressed air fed at an effectivelysteady rate of about 329 kilograms per minute. The operating pressure inthe reaction vessel overhead gas was steadily about 0.41 megapascalgauge. The reaction vessel was operated in a substantially adiabaticmanner so that the heat of reaction elevated the temperature of theincoming feeds and evaporated much of the incoming solvent. Measurednear the mid-elevation of the reaction medium, the operating temperaturewas about 153.3° C.

Reaction medium was withdrawn from the side of the reaction vessel at anelevation of about 14 meters through a wall hole that had an insidecircular diameter of about 0.076 meters. The withdrawn reaction mediumwas conveyed through a substantially horizontal conduit made of nominal3-inch Schedule 10S piping components into the side of a substantiallyvertical external de-aeration vessel. The external de-aeration vesselhad an inside circular diameter of about 0.315 meters, being constructedprimarily of nominal 12-inch Schedule 10S pipe. The horizontalcross-sectional area inside the external de-aeration vessel was thusabout 0.0779 meters squared. This compares to the horizontalcross-sectional area inside the reaction vessel of about 4.67 meterssquared for the elevation where the reaction medium was withdrawn. Thus,the ratio of the smaller to the greater horizontal cross-sectional areawas about 0.017.

The external de-aeration vessel extended downwards from the elevation ofentering reaction medium by about 1.52 meters before transitioning downin diameter to match a bottom outlet flow conduit. An effectively steadyflow rate of about 433 kilograms per minute of substantially de-aeratedslurry comprising crude terephthalic acid exited from the bottom of theexternal de-aeration vessel. Thus, the substantially de-aerated slurryin lower elevations of the nominal 12-inch de-aeration vessel had adownwards superficial velocity that was about 0.093 meters per second;and there was not a deleterious entrainment of oxidant in this exitingslurry. The exiting slurry was conveyed forward by a flow conduit madeof nominal 3-inch Schedule 10S piping components to connect withdownstream processing equipment. In this example, the means forcontrolling the flow rate of withdrawn reaction medium was located inthe flow exiting the bottom of the de-aeration vessel, though othercontrol locations are possible and useful.

The external de-aeration vessel extended above the elevation at whichreaction medium entered by about 14 meters before transitioning from anominal 12-inch piping size down in diameter to match an upper outletflow conduit made of nominal 2-inch Schedule 10S piping components. Theseparated off-gas from the external de-aeration vessel was conveyedthrough this nominal 2-inch conduit to join the main flow of off-gasleaving the top of the reaction vessel.

In this example, the ratio of the production rate of undesirable of4,4′-dicarboxystilbene to the production rate of desired terephthalicacid was measured by HPLC-MS on three separate samples of slurry productas about 2.3, 2.7, and 3.2 ppmw-averaging about 2.7 ppmw. This isimportantly reduced compared to Example 5. The concentration ofpara-xylene in the liquid phase of the slurry exiting from the elevatedside outlet was measured by calibrated GC on three separate samples ofslurry product as about 86, 87 and 91 ppmw-averaging about 88 ppmw. Theconcentration of para-tolualdehyde in the liquid phase of the exitingslurry was measured by calibrated GC on said separate samples of slurryproduct as about 467, 442, and 423 ppmw-averaging about 444 ppmw. Thisis a conversion and yield improvement in the withdrawn slurry flowcompared to Example 5.

EXAMPLES 7-10

Examples 7-10 are calculated examples relating particularly to theinitial dispersion of para-xylene into the reaction medium, but alsodemonstrating other aspects of the present invention.

EXAMPLE 7

This example relates to feeding of vaporized para-xylene. In thiscalculated example, para-xylene feed is heated and vaporized beforeadmission to the reaction medium. This aids initial dispersion of thepara-xylene. It provides enlarged entering volumes and facilitatesincreased velocities. Furthermore, it retards the transfer of theincoming para-xylene into the bulk liquid phase and causes thepara-xylene feed to move toward the reactive liquid phase in betterharmony with the gaseous feeding of molecular oxygen.

In this example, a bubble column oxidizer vessel has a vertical,cylindrical body with an inside diameter of 2.44 meters. The height ofthe bubble column oxidizer vessel is 32 meters from lower tangent line(TL) to upper TL. The vessel is fitted with 2:1 elliptical heads at thetop and bottom of the cylinder. The operating level is about 25 metersof reaction medium above the lower TL. The feed of filtrate solvent,which is separated from para-xylene, enters at a rate of 18.4 kilogramsper second through a 0.076 meter circular diameter entry hole throughthe reaction vessel wall at an elevation of 4.35 meters above the lowerTL. The feed rate of reflux solvent is about 14.3 kilograms per secondinto the gas-disengaging zone above the operating level of the reactionmedium. The feed rate of compressed air is about 9 kilograms per secondthrough an oxidant sparger essentially the same as in Examples 4 and 6.Slurry containing about 31 weight percent solids is withdrawn from thereaction medium through a side draw leg essentially the same as inExamples 4 and 6. The pressure in the headspace above the reactionmedium is about 0.50 megapascal gauge. The contents of water and ofcobalt, bromine and manganese within the liquid portion of the reactionmedium are essentially the same as in Example 4.

The feed rate of para-xylene is 1.84 kilograms per second. Prior torelease into the reaction medium, the feed stream of liquid-phasepara-xylene is pressurized and then vaporized at a pressure of about0.69 megapascal gauge by heating from a storage temperature of about 40°C. up to a temperature of about 233° C. This requires about 1.3megajoules per second of heat input to the feed stream of para-xylene. Aheat exchanger utilizing steam at 4 megapascal is employed for thisduty, but any other energy source of sufficient temperature will sufficeequally, including waste heat from process fluids. This heat inputrepresents about 5 percent of the heat of reaction for para-xyleneconversion to terephthalic acid. Removal of this additional heat loadcauses the reaction medium temperature to rise somewhat at constantpressure, in comparison to feeding para-xylene liquid. (See Example 8.)The temperature is about 162° C. measured near the mid-elevation of thereaction medium. Optionally, pressure could be lowered to reducereaction temperature to 160° C. measured near the mid-elevation of thereaction medium.

The volumetric flow of vaporized para-xylene is about 0.084 cubic metersper second. This flow is admitted to the reaction vessel at an elevationof 0.1 meters above the lower TL of the vessel through 3 conduitsconnected in parallel. Adjacent to the reaction vessel, each conduit ismade from nominal 1.5-inch piping components and connects to a circularhole of equal diameter in the vessel wall. The 3 wall holes are situatedwith 120-degree horizontal, azimuthal spacing from each other. Thesuperficial velocity of each entering stream of para-xylene isapproximately 21 meters per second, and the entering para-xylene isbeing dispersed within the reaction medium at the same time it isdissolving into the reactive liquid phase, where the catalyst speciesprincipally reside.

EXAMPLE 8

This example relates to feeding partly vaporized para-xylene. In thiscalculated example, para-xylene feed is partly vaporized by mixing withthe oxidant supply before admission to the reaction medium. This aidsinitial dispersion of the para-xylene. It provides enlarged enteringvolumes and facilitates increased velocities; and it dilutes theconcentration of para-xylene. Furthermore, it retards the transfer ofthe incoming para-xylene into the bulk liquid phase and causes thepara-xylene feed to move toward the reactive liquid phase in betterharmony with the gaseous feeding of molecular oxygen.

In this example, a bubble column oxidizer vessel has a vertical,cylindrical body with an inside diameter of 2.44 meters. The height ofthe bubble column oxidizer vessel is 32 meters from lower tangent line(TL) to upper TL. The vessel is fitted with 2:1 elliptical heads at thetop and bottom of the cylinder. The operating level is about 25 metersof reaction medium above the lower TL. The feed of filtrate solvent,which is separated from para-xylene, enters at a rate of 18.4 kilogramsper second through a 0.076-meter circular diameter entry hole throughthe reaction vessel wall at an elevation of 4.35 meters above the lowerTL. The feed rate of reflux solvent into the gas-disengaging zone abovethe operating level of the reaction medium is about 12.8 kilograms persecond. The feed rate of compressed air is about 9 kilograms per secondthrough an oxidant sparger similar to the one in Examples 4 and 6, butmodified as noted below. Slurry containing about 31 weight percentsolids is withdrawn from the reaction medium through a side draw legessentially the same as in Examples 4 and 6. The pressure in theheadspace above the reaction medium is about 0.50 megapascal gauge. Thecontents of water and of cobalt, bromine and manganese within the liquidportion of the reaction medium are essentially the same as in Example 4.

The feed rate of para-xylene is again 1.84 kilograms per second. Thisflows as a liquid through conduits to the interior of the oxidantsparger where the liquid is released into the compressed air at 4positions using liquid spray nozzles, as known in the art. Optionally,open ended liquid conduits or gas-liquid spray nozzles may be employedat the point where liquid is admitted to the oxidant sparger. As asafety precaution, 4 temperature sensors are placed within the oxidantsparger. These temperature sensors are connected to alarms andinterlocks to shut off the supply of oxidant and para-xylene if hightemperatures are detected. With the compressed air supply at about 80°C., owing to the heat of compression without an aftercooler on the finalcompression stage, and with the feed para-xylene at about 40° C.,approximately 17 weight percent of the para-xylene is vaporized at thepressure prevailing inside the oxidant sparger. The remaining liquidpara-xylene is carried into the reaction medium with the gas in twophase flow commingled with the gas at velocities approaching those ofthe gas flow. In addition, said remaining liquid helps flush from theoxidant sparger any solids that have intruded, according to aspects ofthe invention.

The temperature is about 160° C. measured near the mid-elevation of thereaction medium. Since no additional energy has been added to any feedstream, this is about the same as Examples 4 and 6.

Optionally, either the compressed air feed or the para-xylene feed canbe pre-heated before mixing in the oxidant sparger in order to increasethe fraction of para-xylene that enters the reaction medium as vapor.For example, a heat input of 300 kilojoules per second to thepara-xylene raises its temperature to about 124° C. and increases thefraction of para-xylene flashed to about 33 percent. For example, a heatinput of 600 kilojoules per second to the compressed air raises itstemperature to about 146° C. and increases the fraction of para-xyleneflashed to about 54 percent. In both cases lower grade energy isrequired for heating than in Example 7. In fact, the waste heat from theoff-gas from the reaction medium can be used as all or part of the heatsource. However, when an amount of energy is added to the feeds, thetemperature of the reaction medium will rise slightly, settling, at thestated pressure, flows and phase compositions, between 160° C. and 162°C. measured near the mid elevation. Optionally, the pressure can beadjusted to adjust temperature. In addition, when an amount of energy isadded to the feeds, the amount of solvent fed to the reaction vessel isadjusted when it is desired to hold solids fraction approximatelyconstant. For example, the reflux solvent flow varies between about 12.8and about 14.3 kilograms per second in Examples 7 and 8, depending onthe amount of energy added, in order to hold solids approximatelyconstant near 31 weight percent.

EXAMPLE 9

This example relates to feeding para-xylene away from the wall of thereaction vessel using a liquid eductor. In this calculated example,initial dispersion of para-xylene liquid feed is improved by using aneductor employing liquid flow as the motive force. The reactor of thisexample has the same mechanical configuration and process boundaryconditions as Example 4 with the exceptions described below. Thecommingled liquid-phase stream of para-xylene plus filtrate solvententers through the reaction vessel wall at the same elevation throughthe same nominal 3-inch flow conduit. However, rather than the internalliquid-phase feed distribution system of Example 4, the commingledliquid-phase feed is released into the reaction medium as the motivefluid in flow eductor as known in the art and as shown in the diagram ofFIG. 26. The eductor is designed for a pressure difference of 0.1megapascal on the motive fluid. The eductor is located and oriented withthe flow plume exiting vertically upwards along the axial center line ofthe reaction vessel at an elevation about 4.5 meters above the lower TL.The volume of reaction medium educted and commingled with the motiveliquid varies with time depending upon stochastic bubble swarm events inthe bubble column at the eduction inlet. However, the time averagededucted flow is greater than the motive fluid flow thus providing a morerapid dilution of incoming para-xylene. Subsequent mixing and chemicalreaction occurs according to the usual stochastic events in the bubblecolumn.

EXAMPLE 10

This example relates to feeding para-xylene away from the wall of thereaction vessel using a gas and liquid eductor. In this calculatedexample, initial dispersion of para-xylene feed is improved by using aneductor employing gas flow as the motive force. The reactor of thisexample has the same mechanical configuration and process boundaryconditions as Example 4, with the exceptions described below. Theoctagonal oxidant sparger and the liquid-phase feed distribution systemare both removed. Instead, the incoming oxidant stream and thecommingled liquid-phase feed of para-xylene plus filtrate solvent areconveyed though separate conduits to the interior of the reactionvessel. There, both streams are combined as motive fluids at the inletof a flow eductor as known in the art and as shown in the diagram ofFIG. 27. The eductor is aligned vertically along the axial centerline ofthe reaction vessel. It is positioned with outlet facing downward andlocated 0.2 meters below the lower tangent line of the reaction vessel.The eductor is designed for a pressure difference of 0.1 megapascal onthe motive fluids. Two temperature sensors are located near where thecompressed air and para-xylene feeds first combine. These temperaturesensors are connected to alarms and interlocks to shut off the supply ofoxidant and para-xylene if high temperatures are detected.

The volume of reaction medium educted is increased compared Example 9and the initial dilution of incoming para-xylene is further improved. Inaddition, the liquid phase portion of the reaction medium with highestlocal concentrations of para-xylene is even more directly staged againstthe gas-phase portion with highest concentration of molecular oxygen.Subsequent mixing and chemical reaction occurs according to the usualstochastic events in the bubble column.

EXAMPLES 11-13

Examples 11-13 are calculated examples relating particularly to usingflows of liquid from the reaction medium in conduits to aid the initialdispersion of para-xylene into the reaction medium, but alsodemonstrating other aspects of the present invention.

EXAMPLE 11

This example relates to using a flow conduit within the reaction vesselto transport liquid to aid the initial dispersion of enteringpara-xylene. The reactor of this example has the same mechanicalconfiguration and process boundary conditions as Example 4, with theexceptions described below. Reference is made to the diagram of FIG. 24.The commingled liquid-phase stream of para-xylene plus filtrate solvententers through the reaction vessel wall through a nominal 3-inch flowconduit similar to Example 4. However, the internal liquid-phase feeddistribution system of Example 4 is removed and said commingled liquidflow is instead released into a flow conduit. The flow conduit has acircular inside diameter of about 0.15 meters for most of its length,including its lower terminus, which is 1 meter above the lower TL of thevessel. The flow conduit rises vertically to a total height of 21 metersabove the lower TL of the vessel. At a height of 20 meters above thelower TL of the vessel, the flow conduit expands to have an inside crosssectional area of 0.5 square meters while rising in height for another 1meter. This upper, larger diameter section of said flow conduit may beconceived as an internal de-aeration vessel, and it is actually formedin part using the wall of the reaction vessel. The entirety of the flowconduit is located within the reaction vessel. At the top inlet to theflow conduit, the reaction medium is greatly depleted of para-xylene andpara-tolualdehyde, though important concentrations of para-toluic acidand 4-carboxybenzaldehyde exist. Reaction medium entering the top ofsaid flow conduit substantially de-aerates, creating a denser medium onthe inside of said flow conduit than in the rest of the reaction vessel.The slurry within the flow conduit moves downward at a rate estimated tobe about 150 kilograms per second, at which point the flowing pressuredrop, integrated over the entire length of said flow conduit, comes intobalance with the density difference between inside and outside,integrated over the entire length of said flow conduit. Of thisdownwards flow of slurry, about 104 kilograms per second is liquid,amounting to about 69 weight percent. The feed flow of intimatelycommingled para-xylene and filtrate solvent, totaling about 20.2kilograms per second, is admitted to the said flow conduit about 5meters above the lower TL. This mixture then travels down the flowconduit an additional 4 meters, about 27 conduit diameters, in less than1 second and becomes appreciably mixed. The concentration of para-xyleneis thus usefully reduced to about 15,000 ppmw before being released intothe main body of reaction medium in the bubble column. Subsequent mixingand chemical reaction occurs according to the usual stochastic events inthe bubble column.

EXAMPLE 12

This example relates to using a flow conduit external to the reactionvessel to transport liquid to aid the initial dispersion of enteringpara-xylene. The reactor of this example has the same mechanicalconfiguration and process boundary conditions as Example 11 with theexceptions described below and with reference to the diagram of FIG. 25.The internal flow conduit is removed and replaced with an external flowconduit. The section of conduit connecting the reaction vessel to theexternal de-aeration section has an inside circular diameter of 0.30meters and is located 20 meters above the lower TL. The inside circulardiameter of the external de-aeration section is 1 meter and its heightis 2 meters. The inside circular diameter of the flow conduit below thede-aeration section is 0.20 meters allowing for larger flows using aboutthe same available elevation head. A flow sensor and a flow controlvalve are included with the flow conduit in order to control the flowrate in the desired range. For example, the flow control is set to allow150 kilograms per second of slurry transport, the same as is estimatedto occur via the internal flow conduit of Example 11. The commingledliquid-phase stream of para-xylene and filtrate solvent is admitted tothe external flow conduit about 5 meters above the lower TL of thereaction vessel. The outlet of the external flow conduit connects to thebottom head of the reaction vessel. Thus, the concentration ofpara-xylene is again usefully reduced to about 15,000 ppmw before beingreleased into the main body of reaction medium in the bubble column.Subsequent mixing and chemical reaction occurs according to the usualstochastic events in the bubble column. The product slurry withdrawalfor post-processing is via a branch from said flow conduit below thede-aeration section and above the addition of the liquid-phase stream ofpara-xylene and filtrate solvent, thus avoiding the need for a separatesystem for removing and de-aerating slurry.

EXAMPLE 13

This example relates to using a flow conduit comprised of sections bothexternal and internal to the reaction vessel to transport liquid to aidthe initial dispersion of entering para-xylene. This calculated exampleis identical to Example 12 except that a second branch in the externalflow conduit is located about 3 meters above the lower TL of thereaction vessel, which is below the addition point of commingledliquid-phase stream of para-xylene and filtrate solvent. The secondbranch flow conduit also has an inside circular diameter of 0.20 meters.A separate flow control valve is placed in the second branch flowconduit, again to regulate the flow. The branch flow conduit penetratesthrough the side wall of the reaction vessel 3 meters above the lowerTL, and the branch flow conduit extends inside the wall of the reactionvessel by 0.4 meters. Thus, the branch conduit comprises sections bothexternal and internal to the reaction vessel. Flow may be admitted tothe reaction vessel through either or both of the bottom-head conduitexit or the side-wall-internal conduit exit and in any ratio.

The invention has been described in detail with particular reference topreferred embodiments thereof, but will be understood that variationsand modification can be effected within the spirit and scope of theinvention.

1. A process comprising: (a) forming a three-phase medium in a reactionzone of a bubble column reactor; (b) introducing at least a portion ofsaid three-phase medium into a deaeration vessel having one or moregenerally upright sidewalls defining a deaeration zone, wherein themaximum horizontal cross-sectional area of said deaeration zone is lessthan about 25 percent of the maximum horizontal cross-sectional area ofsaid reaction zone; and (c) separating at least a portion of the gasphase of said three-phase medium from the solid and liquid phases ofsaid medium in said deaeration zone to thereby form a substantiallydeaerated slurry, wherein said separating is caused primarily by thenatural buoyancy of said gas phase in said solid and liquid phases. 2.The process of claim 1 wherein the maximum horizontal cross-sectionalarea of said deaeration zone is in the range of from about 0.1 to about10 percent of the maximum horizontal cross-sectional area of saidreaction zone.
 3. The process of claim 1 wherein the volume of saiddeaeration zone is less than about 5 percent of the volume of saidreaction zone.
 4. The process of claim 1 wherein the maximum horizontalcross-sectional area of said deaeration zone is in the range of from0.25 to 4 percent of the maximum horizontal cross-sectional area of saidreaction zone, wherein the volume of said deaeration zone is in therange of from about 0.01 to about 2 percent of the volume of saidreaction zone.
 5. The process of claim 1 wherein the maximum horizontalcross-sectional area of said deaeration zone is less than about 3 meterssquared, wherein the volume of said deaeration zone is less than about 2cubic meters.
 6. The process of claim 5 wherein said deaeration zone hasa maximum diameter in the range of from about 0.2 to about 2 meters anda maximum height in the range of from about 0.3 to about 5 meters. 7.The process of claim 1 wherein said substantially deaerated slurrycomprises less that about 5 volume percent gas.
 8. The process of claim1 wherein said substantially deaerated slurry comprises less that about2 volume percent gas.
 9. The process of claim 1 wherein said solid andliquid phases flow downwardly in said deaeration zone at an averagevelocity that is not more than about 75 percent of the natural risevelocity of 0.5 centimeter bubbles of said gas phase in said solid andliquid phases.
 10. The process of claim 1 wherein said solid and liquidphases flow downwardly in said deaeration zone at maximum velocity lessthan about 50 centimeters per second.
 11. The process of claim 1 whereinsaid deaeration zone has a substantially constant horizontalcross-sectional area.
 12. The process of claim 11 wherein saiddeaeration zone is substantially cylindrical in shape.
 13. The processof claim 1 wherein said reaction zone has a maximum length (L), amaximum diameter (D), and an L:D ratio in the range of from about 6:1 toabout 30:1.
 14. The process of claim 1 wherein at least a portion ofsaid deaeration zone is located in said bubble column reactor.
 15. Theprocess of claim 1 wherein all of said deaeration zone is locatedoutside of said bubble column reactor.
 16. The process of claim 1wherein at least a portion of the separated gas phase is re-introducedinto said reaction zone.
 17. The process of claim 1 wherein said processfurther comprises oxidizing an oxidizable compound in the liquid phaseof said reaction medium while said reaction medium is in said reactionzone.
 18. The process of claim 17 wherein said oxidizing is carried outin a manner such that when said reaction zone is theoreticallypartitioned into 2,000 discrete horizontal slices of equal volume, lessthan 40 of said horizontal slices have a gas hold-up less than 0.3 on atime-averaged and volume-averaged basis.
 19. The process of claim 18wherein the entire volume of said reaction medium has a gas hold-up ofat least about 0.4 on a time-averaged and volume-averaged basis.
 20. Theprocess of claim 17 wherein said oxidizable compound is an aromaticcompound.
 21. The process of claim 17 wherein said oxidizable compoundis selected from the group consisting of para-xylene, meta-xylene,para-tolualdehyde, meta-tolualdehyde, para-toluic acid, meta-toluicacid, acetaldehyde, and combinations of two or more thereof.
 22. Theprocess of claim 17 wherein said oxidizable compound is para-xylene. 23.The process of claim 17 wherein said oxidizing causes the formation ofsolids in said reaction medium.
 24. The process of claim 17 wherein saidoxidizing causes at least about 10 weight percent of said oxidizablecompound to form solids in said reaction medium.
 25. The process ofclaim 17 wherein said reaction medium comprises in the range of fromabout 5 to about 40 weight percent solids.
 26. The process of claim 17wherein said oxidizing is carried out in the presence of a catalystsystem comprising cobalt.
 27. The process of claim 26 wherein saidcatalyst system further comprises bromine and manganese.
 28. The processof claim 17 wherein said process further comprises removing at least aportion of said reaction medium from said reaction zone via an elevatedopening located above the lower end of said reaction medium.
 29. Theprocess of claim 28 wherein said process further comprises introducingthe removed reaction medium into said deaeration vessel.
 30. The processof claim 29 wherein said reaction zone has a maximum diameter (D),wherein said elevated opening is located at least about ID from thelower end of said reaction zone.
 31. A process comprising: (a) oxidizingan oxidizable compound in a liquid phase of a three-phase reactionmedium contained in a bubble column reactor, wherein said bubble columnreactor includes a vessel shell defining an outlet; (b) introducing atleast a portion of said three-phase reaction medium into a deaerationvessel comprising one or more generally upright sidewalls at leastpartly disposed within said vessel shell, wherein said upright sidewallsdefine an internal deaeration zone that is in fluid flow communicationwith said outlet; and (c) separating at least a substantial portion ofthe gas phase of said three-phase reaction medium from the solid andliquid phases of said reaction medium in said deaeration zone to therebyform a substantially deaerated slurry, wherein said separating is causedprimarily by the natural buoyancy of said gas phase in said solid andliquid phases.
 32. The process of claim 31 wherein the maximumhorizontal cross-sectional area of said deaeration zone is less thanabout 25 percent of the maximum horizontal cross-sectional area of saidreaction zone.
 33. The process of claim 31 wherein the maximumhorizontal cross-sectional area of said deaeration zone is in the rangeof from about 0.1 to about 10 percent of the maximum horizontalcross-sectional area of said reaction zone.
 34. The process of claim 31wherein the volume of said deaeration zone is less than about 5 percentof the volume of said reaction zone.
 35. The process of claim 31 whereinsaid deaeration vessel presents an open upper end for receiving saidreaction medium from said reaction zone.
 36. The process of claim 31wherein said substantially deaerated slurry comprises less that about 5volume percent gas.
 37. The process of claim 31 wherein said deaerationzone is substantially cylindrical in shape.
 38. The process of claim 31wherein said oxidizable compound is para-xylene.
 39. The process ofclaim 31 wherein said oxidizing causes at least about 10 weight percentof said oxidizable compound to form solids in said reaction medium. 40.A process comprising: (a) oxidizing an oxidizable compound in a liquidphase of a three-phase reaction medium contained in a reaction zone of abubble column reactor, wherein said bubble column reactor includes avessel sidewall defining a side outlet that is elevated from the lowerend of said reaction zone; (b) withdrawing at least a portion of saidthree-phase reaction medium from said reaction zone via said sideoutlet; (c) introducing at least a portion of the withdrawn three-phasereaction medium into a deaeration vessel comprising one or moregenerally upright sidewalls defining a deaeration zone; and (d)separating at least a substantial portion of the gas phase of saidthree-phase reaction medium from the solid and liquid phases of saidreaction medium in said deaeration zone to thereby form a substantiallydeaerated slurry, wherein said separating is caused primarily by thenatural buoyancy of said gas phase in said solid and liquid phases. 41.The process of claim 40 wherein the maximum horizontal cross-sectionalarea of said deaeration zone is less than about 25 percent of themaximum horizontal cross-sectional area of said reaction zone.
 42. Theprocess of claim 40 wherein the maximum horizontal cross-sectional areaof said deaeration zone is in the range of from about 0.1 to about 10percent of the maximum horizontal cross-sectional area of said reactionzone.
 43. The process of claim 40 wherein the volume of said deaerationzone is less than about 5 percent of the volume of said reaction zone.44. The process of claim 40 wherein said substantially deaerated slurrycomprises less that about 5 volume percent gas.
 45. The process of claim40 wherein said deaeration zone is substantially cylindrical in shape.46. The process of claim 40 wherein said oxidizable compound ispara-xylene.
 47. The process of claim 40 wherein said oxidizing causesat least about 10 weight percent of said oxidizable compound to formsolids in said reaction medium.
 48. The process of claim 40 wherein saidreaction zone has a maximum diameter (D), wherein said side opening islocated at least about ID from the lower end of said reaction medium.49. A process for making purified terephthalic acid, said processcomprising: (a) oxidizing para-xylene in a liquid phase of a three-phasereaction medium contained in a reaction zone of a bubble column reactorto thereby form crude terephthalic acid; (b) introducing at least aportion of said three-phase reaction medium into a deaeration vesselhaving one or more generally upright sidewalls defining a deaerationzone, wherein the maximum horizontal cross-sectional area of saiddeaeration zone is less than about 25 percent of the maximum horizontalcross-sectional area of said reaction zone; (c) separating at least asubstantial portion of the gas phase of said three-phase medium from thesolid and liquid phases of said medium in said deaeration zone tothereby form a substantially deaerated slurry comprising said crudeterephthalic acid, wherein said separating is caused primarily by thenatural buoyancy of said gas phase in said solid and liquid phases; and(d) oxidizing at least a portion of said crude terephthalic acid fromsaid deaerated slurry to thereby form purer terephthalic acid.
 50. Theprocess of claim 49 wherein the maximum horizontal cross-sectional areaof said deaeration zone is in the range of from about 0.1 to about 10percent of the maximum horizontal cross-sectional area of said reactionzone.
 51. The process of claim 50 wherein said substantially deaeratedslurry comprises less that about 5 volume percent gas.
 52. The processof claim 49 wherein said oxidizing in said secondary oxidation reactordecreases the concentration of 4-CBA in said crude terephthalic acid byat least about 200 ppmw to thereby form said purer terephthalic acid,wherein said crude terephthalic acid has an average 4-CBA concentrationof at least about 400 ppmw and said purer terephthalic acid has anaverage 4-CBA concentration of less than about 400 ppmw.
 53. The processof claim 49 wherein said oxidizing in said secondary oxidation reactordecreases the average concentration of 4-CBA in said crude terephthalicacid by at least about 400 ppmw to thereby form said purer terephthalicacid, wherein said crude terephthalic acid has an average 4-CBAconcentration of at least about 800 ppmw and said purer terephthalicacid has an average 4-CBA concentration of less than about 250 ppmw. 54.The process of claim 49 wherein said oxidizing in said secondaryoxidation reactor is carried out at an average temperature at leastabout 10° C. greater than the average temperature of said oxidizing insaid bubble column reactor, wherein said oxidizing in said bubble columnreactor is carried out at an average temperature in the range of fromabout 125 to about 200° C., wherein said oxidizing in said secondaryoxidation reactor is carried out at an average temperature in the rangeof from about 160 to about 240° C.
 55. The process of claim 49 whereinsaid oxidizing in said secondary oxidation reactor is carried out at anaverage temperature in the range of from about 20 to about 80° C.greater than the average temperature of said oxidizing in said bubblecolumn reactor, wherein said oxidizing in said bubble column reactor iscarried out at an average temperature in the range of from about 140 toabout 180° C., wherein said oxidizing in said secondary oxidationreactor is carried out at an average temperature in the range of fromabout 180 to about 220° C.
 56. The process of claim 49 wherein asubstantial portion of said crude terephthalic acid exists as solidcrude terephthalic acid particles having an average BET surface area ofat least about 0.6 meters squared per gram.
 57. The process of claim 56wherein said solid crude terephthalic acid particles have an averageparticle size in the range of from about 20 to about 150 microns. 58.The process of claim 57 wherein a substantial portion of said solidcrude terephthalic acid particles are formed of a plurality ofagglomerated sub-particles having an average particle size in the rangeof from about 0.5 to about 30 microns.
 59. The process of claim 58wherein said solid crude terephthalic acid particles have an averageparticle size in the range of from about 30 to about 120 microns,wherein said sub-particles have an average particle size in the range offrom about 1 to about 15 microns.
 60. The process of claim 49 whereinsaid deaerated slurry comprises a mother liquor and said crudeterephthalic acid, wherein said process further comprises replacing atleast 50 weight percent of said mother liquor in said deaerated slurrywith a replacement solvent to thereby provide a solvent-exchanged slurrycomprising said replacement solvent and said crude terephthalic acid,wherein said process further comprises introducing saidsolvent-exchanged slurry into said secondary oxidation reactor.
 61. In abubble column reactor for contacting a predominately liquid-phase streamand a predominately gas-phase stream, the improvement comprising: avessel shell defining a reaction zone and a vessel outlet; and adeaeration vessel coupled in fluid flow communication with said vesseloutlet, wherein said deaeration vessel presents an upper inlet and alower outlet, wherein said deaeration vessel comprises one or moreupright sidewalls defining an internal deaeration zone in fluid flowcommunication with said upper inlet and said lower outlet, wherein themaximum horizontal cross-sectional area of said deaeration zone is lessthan about 25 percent of the maximum horizontal cross-sectional area ofsaid reaction zone.
 62. The bubble column reactor of claim 61 whereinthe maximum horizontal cross-sectional area of said deaeration zone isin the range of from about 0.1 to about 10 percent of the maximumhorizontal cross-sectional area of said reaction zone.
 63. The bubblecolumn reactor of claim 61 wherein the volume of said deaeration zone isless than about 5 percent of the volume of said reaction zone.
 64. Thebubble column reactor of claim 61 wherein said deaeration zone has amaximum diameter in the range of from about 0.2 to about 2 meters and amaximum height in the range of from about 0.3 to about 5 meters.
 65. Thebubble column reactor of claim 61 wherein said deaeration zone has asubstantially constant horizontal cross-sectional area.
 66. The bubblecolumn reactor of claim 65 wherein said deaeration zone is substantiallycylindrical in shape.
 67. The bubble column reactor of claim 61 whereinsaid reaction zone has a maximum length (L), a maximum diameter (D), andan L:D ratio in the range of from about 6:1 to about 30:1.
 68. Thebubble column reactor of claim 67 wherein said L is in the range of fromabout 20 to about 75 meters, said D is in the range of from about 2 toabout 10 meters, and said L:D is in the range of from about 8:1 to about20:1.
 69. The bubble column reactor of claim 67 wherein said vessel saidreaction zone presents a normally-lower end and a normally-upper endspaced from one another by said length (L), wherein said vessel outletis spaced at least about 1D from said normally-lower end of saidreaction zone.
 70. The bubble column reactor of claim 61 wherein atleast a portion of said deaeration zone is located in said vessel shell.71. The bubble column reactor of claim 61 wherein all of said deaerationzone is located outside of said vessel shell.
 72. A system for producingand purifying terephthalic acid, said system comprising: a reactor forfacilitating the production of crude terephthalic acid in a three-phasereaction medium, wherein said reactor defines a reactor outlet fordischarging at least the solid and liquid phases of said reaction mediumfrom said reactor; a deaeration vessel for facilitating thedisengagement of the gas phase of said three-phase reaction medium fromthe solid and liquid phases of said three-phase reaction medium tothereby provide a substantially deaerated slurry, wherein saiddeaeration vessel is coupled in fluid flow communication with saidreactor outlet, wherein said deaeration vessel defines a reaction mediuminlet for receiving said three-phase reaction medium and a slurry outletfor discharging said deaerated slurry, wherein said slurry outlet islocated lower than said reaction medium inlet, wherein said deaerationvessel comprises one or more upright sidewalls defining an internaldeaeration zone vertically disposed between said reaction medium inletand said slurry outlet, wherein said deaeration zone has a maximumheight in the range of from about 0.3 to about 5 meters and a maximumhorizontal cross-sectional area in the range of from about 0.02 to about3 meters squared; and a purification system for receiving said deaeratedslurry from said slurry outlet and purify said deaerated slurry.
 73. Thesystem of claim 72 wherein the maximum horizontal cross-sectional areaof said deaeration zone is in the range of from about 0.1 to about 10percent of the maximum horizontal cross-sectional area of said reactionzone.
 74. The system of claim 72 wherein the volume of said deaerationzone is less than about 5 percent of the volume of said reaction zone.75. The system of claim 72 wherein the volume of said deaeration zone isin the range of from about 0.01 to about 2 percent of the volume of saidreaction zone.
 76. The system of claim 72 wherein the volume of saiddeaeration zone is less than about 2 cubic meters.
 77. The system ofclaim 72 wherein said deaeration zone has a substantially constanthorizontal cross-sectional area.
 78. The system of claim 77 wherein saiddeaeration zone is substantially cylindrical in shape.
 79. The system ofclaim 72 wherein said reaction zone has a maximum length (L), a maximumdiameter (D), and an L:D ratio in the range of from about 6:1 to about30:1.
 80. The system of claim 79 wherein said L is in the range of fromabout 20 to about 75 meters, said D is in the range of from about 2 toabout 10 meters, and said L:D is in the range of from about 8:1 to about20:1.
 81. The system of claim 72 wherein said vessel said reaction zonepresents a normally-lower end and a normally-upper end spaced from oneanother by said length (L), wherein said reactor outlet is spaced atleast about ID from said normally-lower end of said reaction zone. 82.The system of claim 72 wherein at least a portion of said deaerationzone is located in said vessel shell.
 83. The system of claim 72 whereinall of said deaeration zone is located outside of said vessel shell. 84.The system of claim 72 wherein said purification system comprises asecondary oxidation reactor.
 85. A product produced by the process ofclaim
 1. 86. An oxidation product produced by the process of claim 31.87. An oxidation product produced by the process of claim
 40. 88. Aterephthalic acid product produced by the process of claim 49.